memo

Process Proposal for MAb Production in

Dunaliela tertiolecta









Bi Motum

Moving Tomorrow’s Technology





Katie Thompson

Lisa Feenstra

Dieu Pham

Rodolfo Ham-Zhu









Report Submitted: 6/6/2006

For: Dr. Haluk Hamamci

MAb Background and Process Proposal







Therapeutic Protein History



The introduction of monoclonal antibodies has enriched the tapestry of the biological and

biomedical field given that it has extended our understanding of immunology and provided new

strategies for the combating and treatment of immune deficient diseases. Produced through the

use of hybridoma cells for therapeutic purposes, recombinant protein antibodies like Avastin,

which is currently being marketed as a treatment for colon and rectal cancer, have pioneered a

multi-billion dollar biotechnology industry.



Vascular Endothelial Growth Factor, VEGF, a growth signal sent out by the tumor,

enables tumors to receive nutrients. Avastin functions by attaching directly to VEGF therefore

preventing the tumor blood vessels from growing and spreading. Targeting the tumor’s blood

vessels reduces the size of the tumor because in order for tumors to grow they need a constant

supply of oxygen and other nutrients. Thus by cutting off the tumor’s supply line, they can no

longer grow and spread.



Market Considerations________________________________



Antibody drugs are expected to exceed $68 billion in worldwide sales by the year 2010

(Liszewski). The recent prospects of expanding Avastin to treat both lung and breast cancer

patients may play a large part in the therapeutic protein market. Considering that the yearly

number of diagnosed colon cancer patients is about one third of both breast and lung cancer

combined (see Figure 4), and that the dosage for the potential patients is twice that of patients

currently using Avastin, it is no wonder that the demand for this product is expected to clearly

surpass the supply (see Tables 1-3). If we do some rough calculations with current Genentech

capacity in mind, we can approximately quantify the increased demand of this particular product.

As illustrated in Table 2, 1.6 million grams of additional Avastin is needed yearly to meet

demands for all three cancers. Considering that Genentech plans on utilizing a 90,000 liter

facility to meet this demand and current mammalian cell lines produce roughly only 0.3grams of

MAb per liter of production (Hu, Wei-Shou), Table 3 illustrates that the exceeded yearly demand

for this protein alone could be approximately 19,000 grams.





1

Figure 1 – Number of cancer cases annually

Source: Biotechnology Industry Organization



Table 1: Current Avastin Use: Colon Cancer Patients

Avastin sales for 20051 $1,133,000,000

Cost per gram of Avastin1 $5,500

Grams sold in 2005 206,000grams

2

Mg Avastin per kg body weight for colon cancer patients 5mg/Kg

Mg used per dose (using 70kg as average body weight) 350mg

Grams used per year, per patient (assuming one dose every two weeks) 8.4grams

Number of patients in 2005 using Avastin (g sold/g per patient) 24,524

3

Number of colon cancer patients in 2005 154,000

Percentage of colon cancer patients using Avastin 16%



Table 2: Projected Avastin Use Increase: Lung and Breast Cancer

Number of current lung cancer patients3 167975

3

Number of current breast cancer patients 433935

Potential increase in Avastin patients due to lung cancer (using 16% of total) 26876

Potential increase in Avastin patients due to breast cancer (using 16% of 69430

total)

Mg Avastin per Kg body weight for both lung and breast cancer2 10mg/Kg

Mg used per dose (using 70kg as average body weight) 700mg

Grams used per year, per patient (assuming one dose every two weeks) 16.8grams

Yearly increase of Avastin sold due to lung and breast cancer 1.6 million

grams

Yearly Avastin sales increase due to lung and breast cancer $8.9 billion









2

Table 3: Increased Production Demand

Genentech’s expected capacity to meet this demand2 90,000Liters

Yield MAb per liter using current mammalian cell line 0.3grams/Liter

Gram MAb per gram Avastin (approximate)4 .025g/g

Total gram of MAb produced per fermentation 27000 grams

Total gram of Avastin produced per fermentation 1,080,000 grams

Total gram of Avastin needed (current plus potential increase) 1,823,941grams

Total grams of Avastin needed that will exceed Genentech 743,941grams

capacity

Total grams of MAb needed that will exceed Genentech 19,000 grams

capacity

Information sources for Tables 1-3:

1

-Genentech Financial Report 2005

2

-Hu, Jim, Blogs for Industry

3

-American Cancer Society

4

-Food and Drug Administration



In addition to the demand increase due to the existing drugs on the market ( 22% MAb),

the pipeline of 200+ pending therapeutic protein pharmaceuticals (57% MAb) also pushes for

increased production levels (Steiner). Because of the expected rise in required production, most

of the pharmaceutical and biotechnology companies already outsourcing some of their

biopharmaceutical manufacturing are planning to increase their spending in the next two years.

Median contract spending in 2004 was reported as roughly $3 to $4 million per contract

organization per year and this number has been predicted to double by the end of this year (Fox).

The size of the worldwide contract manufacturing market is expected to reach $2.5 billion this

year (Fox). Considering the nearly two-fold increase from the $1.7 billion earned in 2004,

contract manufacturing is looking at a very bright future.





Process Cost Considerations



As competition and production volumes increase in response to the accelerated demands,

incentives for improving efficiencies are expected to take a big leap forward. In other words, low

production costs will be crucial for the success of our company. Although lower production

costs are desired, lower yields are not. When considering current expression systems utilized to

fabricate monoclonal antibodies, there are both advantages and disadvantages in terms of the

bottom line.



There are currently two strategies for producing monoclonal antibodies like Avastin:

CHO cells lines and plant cultures. There are many differences between plant, and CHO based

expression systems. These differences become very important when considering how to make

our organization as competitive as possible and are summarized below in Table 4.







3

Table 4: Comparison of CHO and Plant Cell Expression Systems



Expression CHO Plant Cells (Traditional and algae)

System



Shear Resistance Lowest Medium

Respiration Rate Lowest Medium

Growth Rate Medium Slow (traditional plants such as tobacco)

Fast (algae)

Doubling Time 1 day One week (traditional plant)

3.5 hours (algae)

Excreted Protein Yes Yes

Post-translational Yes. Post-translational Yes, with some restrictions.

Modifications modifications are almost Plant proteins may be immunogenic.

identical to humans.

Problems faced in Oxygen transfer is often a Getting enough light to cells is often a

production problem because cells cannot problem. Plant cultures can grow densely

withstand high agitation or and become viscous.

aeration.

Media Complex. Contains serum, Simple.

which is expensive and easily

contaminated.

Safety Can transmit disease-causing Cannot transmit animal diseases.

agents such as viruses.

Other Suitable for large and Easy to scale up.

complicated proteins.





Currently monoclonal antibodies like Avastin (Bevacizumab) are being mass produced in

a Chinese Hamster Ovary (CHO) mammalian cell expression system in a nutrient medium

containing the antibiotic gentamicin. Through post-translational modifications and

Bevacizumab’s covalent structure, the molecule was modified to posses similar characteristics of

a human IgG1 (European Medicines Agency).1 The current expression vector used to express

this antibody is the plasmid pSVID5.ID.LLnspeV.xvegf36HC.LC. The plasmid is encoded with

Bevacizumab and then inserted into CHO DP-12 (European Medicines Agency)2. To begin the

inoculation train 20L of culture cells are inoculated, and the cells sub-cultivated every three to

four days in a subsequent series of increasing volumes (80L  400L 2000 L). An aliquot of

the cells are then placed in a 12000L CHO cell fed-batch suspension culture for harvesting

(European Medicines Agency).3 Post fermentation, the cells are run through a recovery train

which isolates and purifies the therapeutic protein. Due to the animal protein requirements of this





1

Scientific Discussion. European Medicines Agency. 4 Apr. 2006. pg 7.

2

IBIDEM. pg 4.

3

IBID. pg 5.





4

cell line, viral contamination is an issue and involves rather expensive filtering and deactivation

procedures. Overall, the current technology is expensive and time consuming.



Alternatively, two approaches exist to produce these proteins from plants: plant cell and

algae suspension culture. Due to damage by twisting motion (in turbulent mix regime) to plant

cells, gas-to-liquid mass transfer is limited. Because of plant’s low cell growth rate, aseptic

conditions in plant suspension cultures have to be maintained for two to four weeks. Although

shear stress affects the growth of algaein a concentric tube airlift photobioreactor (Contreras, et.

al.), the effects are less drastic than in plant cell cultures (Wu, et. al). In addition, algae grows

much faster than plant cells. For these reasons, algae suspension cultured in photobioreactors are

the optimum choice. Photon flux density (PFD) is one of the most important limiting and to-be-

controlled variables (Merchuck, et. al.). Mathematical models that relate biomass productivity to

average photon flux density –which can be controlled by varying agitation speed, have been

developed by Merchuck. According to Merchuck’s model, the optimum algae concentration is

108 cell/ml – higher than in plant cells. Assuming that productivity is proportionally related to

cell concentration, algae cultures should yield much higher production rates.



Algae is a good candidate because of substantially lowered manufacturing and

maintenance costs as well as higher yields. In order to produce Bevacizumab, an expression

system with post translational modification (PTM) is usually required. PTM’s such as

phosphorylation, glycosylphosphatidylinositol and ubiquitination have been confirmed to exist in

algae strains such as Dunaliela tertiolecta (Kwon et al., 2005). Like CHO cells, algae cell

cultures need to be first grown in a laboratory. But as pointed out in Table 5 below, the algae

cells grow much more quickly than CHO cells. The algae strain Dunaliela tertiolecta will first be

inoculated in disposable wave bags with magnetic stirrer bars for light and CO2 mixing. Wave

bags are superior for inoculation usage because they are biocompatible and have temperatures

tolerance from 0 to 50ºC. In addition, given that wave bags are disposable, a separate CIP unit is

not required for this step of the process. This allows for easier validation and a reduced factory

start-up time. Once the suspension culture is made, it can be used to inoculate a

photobioreactor. A fed-batch system can then be used to operate the photobioreactor to grow the

algae cells and produce the desired product. After the initial inoculation, a sample of the

previous batch culture can be saved to inoculate the next batch. The media for algae cells does

not have to have precursors and is therefore considered simple and free of viral contaminants.

The algae cells need to have light, carbon dioxide, and nutrient rich water. The media needs to

have a source for carbon, nitrogen, sulfur, and phosphate. Once the growth is complete, the cells

are sent to a recovery train which involves centrifugation for both cell isolation and cell debris

removal, cell lysing to extract the desired product, and further purification steps involving

Protein A columns, Gel Chromatography, and Anion Exchange Chromatography. The general

schematic of the proposed process is represented below in Fig. 2.









5

Inoculation Train Cell Concentration

Upstream Processing Build-up





Growth Photobioreactor Cell Growth



Cell Isolation, Broth

Centrifugation Removal



Cell Lysis





Centrifugation Cell Debris Removal



Downstream Processing

Protein A Column Protein Isolation





Gel Column Improper Protein

Chromatography Removal







Cation Exchange Protein Purification

Column







Anion Exchange Protein Purification

Column







Final Protein Product



Figure 2: Process flowsheet of MAb Production via Dunaliela tertiolecta algal cells









6

The following is an estimated timeline for the completion of the project.



Timeline

Program Phase; 1 Month

1. Facility Program

a. Furnishes all data for plant and decide plant capacity

b. List Operational Description- list of all products to be produced, a general description

of a the production process, expectation of process control, and a statement on the

future plans of the plant

c. List of Organization Chart- anticipated facility staff and their main duties

d. Master List of Required Spaces and why they are needed

e. Functional Resolution Analysis- include warehousing needed and ingredient

preparations

f. Comparative Analysis of Spatial Sizes- compares and contrast areas and volumes

using modular increments

g. Estimation of the volume prediction and area needed

2. Site Analysis

a. Decide on location to ensure that it can accommodate the large facility.

b. Location and property survey

c. Conduct a topographic map of the site and explore the subsurface of the desire area

3. Design Criteria

a. Rough design estimate should be done for each piece of material

4. Utility

a. List all utilities needed and their approximate cost; include also cost of peak demands

for each utility item

b. Report all capacity of all utilities used, ie 50 psig of steam available for heat

exchanger

5. Target data for plant completion and execution

6. Prepare a rough project budget

Conceptual Design Phase; 1.5 Months

1. Process Flow Diagrams

2. Process is scaled up and defined and utilities systems are outline

3. List of equipment and their specifications

4. Facility Layout is developed

5. Decide if project is economically feasible by performing detail economic analysis

6. Develop alternative schemes and estimates

7. Generate a general project schedule

Design Development Phase; 1.5 Months

1. P&ID, instrumentations and flow diagrams

2. Prepare list of all equipment needed

3. Prepare equipment specifications

4. Prepare lists of where equipments can be purchase

5. Prepare Piping diagrams

6. Estimate a general cost by taking vendors quotations

7. Also note equipment that needs to be shipped or transported to plant and their costs







7

8. Identify emergency power requirement

9. Control system Configuration

10. Heating, ventilation, and air conditioning systems are defined if needed

11. Define electrical systems

12. Design layout of plant structure

13. Notify FDA concerning plans for project completion and start-up.

Detail Design Phase; 3-4 Months

1. Prepare detail structural drawing of building

2. Show how construction is performed

3. Prepare detail mechanical drawing of building and operating equipment

4. Prepare detail electrical drawings of all operating equipment

5. Purchase equipment

a. Have manufacturer of all equipment bought pre-tested before shipping

6. Deliverable equipment

a. Backup/Extra Equipment Parts

b. Documentation (include number of copies)-including operating manuals, test

results, guarantees, help manuals

c. Drawings-including final P&IDs and wiring diagrams of all equipment purchased

d. Backup copies of software and important manuals

e. List of company contacts and technical support policy/numbers

7. Construction Bid Documents

8. Construct design of upstream and downstream units

Construction Phase; 4 -5 Months

1. Site Preparations

2. Build plant

3. Include all needed wiring and mechanical parts

4. Install all the needed equipment and check to make sure everything is running properly

Validation Phase; 6 Months

1. Document Preparation- prepare checklist and procedures to be used during filed validation

2. Installation Qualification

3. Operational Qualification



Start-Up









8

Upstream Processing







Culture Requirements





Light Energy



The light will be provided using monochromatic lights at various sizes specified in the

following specification tables. Light transfer is an issue, and becomes increasingly so as density

increases, therefore, the size of the light source increases rather exponentially as the size of the

reactor increases. Chlorophyl absorbs at 650-700 nm; this is equivalent to 4.6x1014-4.3 x1014Hz.

At this energy, Planck’s law indicates 3.3-3.5x1018 photons Joule-1. A proton flux density of

58μmol m-2s-1 has given a maximum biomass yield of 1.2 g mol (Janssen et al., 2000). Assuming

the PFD optimum value to be 58μmol m-2s-1, and taking 3.4 x1018 photons Joule-1

(monochromatic light), the required power per area will be 10.3 Wm-2. Yet a more conservative

PFD estimate of 200-300 μmol m-2s-1(Quigg & García-González) is considered, which would

correspond to 34.46-51.7 W m-2. The light concentrations can be measured using a

photosynthetic photon flux density device (PPFD) using a quantum sensor connected to a

quantum photometer.





Nutrients



To determine the type of media that will be used and the amount of nutrients required,

elements present in the algae strain used for our production line were determined and are shown

below in Table 5. In addition, 3.5% protein/dry weight (w/w) ( Dr. Baez ) is assumed for

protein yield per biomass. This results in a production of 17,150 kg of dry weight per year of

biomass. A summary of the amount of substrate required for the production of such amount of

biomass is shown in Table 6.









9

Table 5: Elements Found in Dunaliela tertiolecta

Element MW Proportion Total Proportion by Normalized

[g/mol] by moles weight weight [%] to 4 g/day

[moles]

C 12 222 2664 64.21 2.57

N 14 38 532 12.82 0.51

P 30.97 1 31 0.75 0.03

S 32 0.28 9 0.22 0.01

K 39 0.36 14 0.34 0.01

Mg 24.3 0.37 9 0.22 0.01

Ca 40 0.019 1 0.02 0.00

Sr 87.62 0.0081 1 0.02 0.00

Fe 55.84 11.3 631 15.21 0.61

Mn 54.9 1.9 104 2.51 0.10

Zn 65.4 1.49 97 2.35 0.09

Cu 63.5 0.67 43 1.03 0.04

Co 58.9 0.01 1 0.01 0.00

Cd 112.4 0.1 11 0.27 0.01

Mo 95.94 0.011 1 0.03 0.00

Total 4148.60806 100 4

Source: Phytoplankton Dynamics Laboratory. University of Texas A&M at Galveston.



Table 6: Substrates Required for 1715kg of Dry Biomass

Compound Substrate Proportion Kg MW Percent Substrate

by weight element element required

[%] needed in [Kg/year]

[Kg/year] substrate

C CO2 64.21 1101.20 44 0.27 4037.7

N KNO3 12.82 219.86 101.1 0.14 1587.7

P K2HPO4 0.75 12.86 136.1 0.23 56.5

S CuSO4·5H2O 0.22 3.77 249.7 0.13 29.3





Ca CaCl2·2H2O 0.02 0.34 147 0.27 1.3

Fe FeCl3, Na2EDTA 15.21 260.85 162.2 0.34 758.2

Mn MnCl2·4H2O 2.51 43.05 147.9 0.37 116.0

Co CoCl2·6H2O 0.01 0.17 237.9 0.25 0.7

Mo NH4)6Mo7O24·4H2O 0.03 0.51 1235.9 0.54 0.9

Trace Various 4.22 72.37 - - -

100 1715

Source: Aizawa and Miyachi, Brown et al., Mil'ko, Grant; Borowitzka and Borowitzka,

McLachlan, Massyuk.



Given that the amount of substrate used in CHO cells does not vary with the production





10

of different proteins (Dr. Stark), it is reasonable to assume that the majority of nutrients are used

for biomass production. From literature, the highest protein yield in plants ranges from 35-50%

protein/dry weight basis. Despite the relatively low substrate use for protein production, it is not

reasonable to ignore the substrates used given that 500 kg is required per year.



Table 7 shows the percent concentration of elements in Bevacizumab and N-

glycosylation. Oxygen is generated from carbon dioxide according to the stoichiometric formula

n CO2 + 2n H2O + ATP + NADPH → (CH2O)n + n O2 + n H2O; for this reason, carbon dioxide

is the limiting reagent. Also, nitrogen and hydrogen percent concentration is higher in the protein

than in N-glycosylation. This is important because mass balances calculations are based on the

protein percent concentration4.



Table 7: Percent Concentration of Elements in Belvacizumab and N-glycosylation

Element Proportion in Proportion in Substrate Kg element Kg substrate

protein N-glycosylation needed needed

(% weight) (% weight) [Kg/yr] [Kg/yr]

C 69.5% 40.8% CO2 347.5 1275

O 12.5% 51.9% CO2 0 0

N 9.25% 1.1% KNO3 46.2 334

H 8.75% 6.2% Various 43.8 Various

Source: US Patent & Trademark Office (for protein amin oacid sequence), Garret & Grisham

(Glycosylation), Julio Baez (for substrate used).







Upstream Design Specifications_________________________







Medium Preparation

The primary method of sterilizing the media is through heat. The gaseous CO2 supply

does not need to be heat sterilized but must be filter sterilized before it enters the reactor. Before

heat sterilizing the media, it should be microfiltered in order to remove any large particles.

Then, the media should be heat sterilized using steam to the point at which the probability of

contamination is very low. Heat sterilizing the media will also degrade some of the media

components. Thus the initial concentration of nutrients in the media must account for this

reduction to the final concentration. After the media is heat sterilized, it must be quarantined

before it is used in order to verify that it is not contaminated.



The liquid media is sterilized in a continuous sterilization system. Continuous

sterilization systems have advantages over batch sterilization systems such as lower energy costs,



4

If percent composition of protein is used in mass balance, nitrogen and hydrogen will be in excess. Thus allowing

protein production at is maximum rate.





11

more efficient use of energy, and the fact that the time to sterilize is lower because the

temperature is higher. The system (detailed specifications listed below in Table 8 ) consists of

two heat exchangers, a retention coil, and a holding tank. Two heat exchangers are used in order

to recover the energy used to heat up the media. The primary heat exchanger has a hot incoming

feed from the sterilized media leaving the retention coil. Energy is recovered when the hot

sterilized media heats up the incoming cold media feed that needs to be sterilized.



The media first enters the primary heat exchanger and is heated from 30˚C to 120˚C. The

cold media is heated up by the hot sterilized media, which is cooled from 130˚C to 50˚C. The

media then enters the secondary heat exchanger where it is further heated by steam to 130˚C.

The media next enters the 33 meters long retention coil at a velocity of 0.11 m/s. It takes the

130˚C media 5 minutes to travel through the retention coil, where the majority of any

contaminants are killed. The sterilized media then returns to the primary heat exchanger, heating

up the fresh, incoming media. The sterilized media is then held in a 10,000 L collection tank. A

detailed schematic of this process is illustrated below in Fig. 3.





Table 8: Media Preparation Specifications

Equipment Material Dimensions

Collection Tank 316L SS 10,000 L. Diameter of 1.62m and

height of 4.86m.

Regen. Heat Exchanger 316L SS 16 plates. Each plate is 40 cm x

30 cm x 11.17 mm.

Steam Heat Exchanger 316L SS 24 plates. Each plate is 40 cm x

30 cm x 11.17 mm.

Retention Coil 316L SS 3 inch inner diameter. 3.5 inch

outer diameter. 91.4 meters long.

Centrifugal Pump 316L SS Magnatex Pumps, ISO 9000

Temperature probe WQ101 Temperature Sensor

pH Probe PHR-212

Flow Meter Celcon PS601C

Liquid Level Meter LV800 Series Level Meter

Acid Supply Tank 316L SS 1,000 L. Diameter of 0.75m and

height of 2.25 m.

Base Supply Tank 316L SS 1,000 L. Diameter of 0.75m and

height of 2.25 m.

Diaphragm Valve High-Purity Two-Way

Diaphragm Valve EW-98613-56

Sampling Probe PFA 100 Probe









12

Temperature

Probe



Incoming CIP



Steam



Incoming Medium Feed Temperature

Probe





Temperature

Start-up Recycle Heat Exchanger Probe







Retention Coil



P-7



Pump Flow Meter

Regenerating Heat Exchanger Steam

Acid Supply Base Supply

Heat Exchanger









Condensate

PLC







PLC Collection

Level Meter

Tank



Sampling Probe





pH Meter









Pump



PLC









Dispense for inoculation medium To bioreactor





Figure 3: Medium Preparation P&ID









13

Inoculation Train

From Dunaliella tertiolecta’s annual biomass production and the saturation

concentration, the volume to process annually should be 343,000 L (sample calculations shown

in Appendix A). This means that 34300 L of inoculate stream is required. Detailed specifications

of the inoculation train are listed below in Table 9.



Growth rate for green algae at light saturation differs from strain to strain. For example,

Scenedesmus obliquus doubles 2.2 times per day, whereas Chrorella pyrenoidosa doubles 9.2

times per day. Since growth rate data is not available for Dunaliela tertiolecta, an estimate has to

be made. Given that most algal strains double at 3.5 hours5, a conservative estimate is 4 hours.

10% inoculants in the inoculation train would, therefore, take 16hours to grow (see calculation in

Appendix A).



Due to absolute sterility as well as cost effectiveness, disposable wave bags were chosen

as the reactor for the inoculation train. Since the largest wave bags produced are 2000L and the

working volume is 50%, the inoculation scheme will be as follows (Fig. 4):



 

10mL 16100mL 161L 1610L 16100L 161000L 16 Photobioreactor

hrs hrs



hrs

 hrs

 hrs

 hrs



Figure 4: Inoculation Flow Sheet



The total time to inoculate the photobioreactor will be 96 hours, or 4 days. Due to the

working capacity of 50%, the actual size of the bags (and test-tube – 10mL) used are two times

that of the loaded volume listed above. The first inoculation will be from a standard Petrie dish

and into a 20mL disposable test tube filled with 9mL of fresh medium using a disposable

inoculation needle. The second inoculation will involve a 200mL wavebag. The bag is first filled

with sterilized air until rigid and then filled with 90mL of fresh medium. The bag is then placed

on the mechanical system and then rocked at 15 rocks per minute at an angle of 6 degrees. The

temperature is allowed to equilibrate and then the inoculant (10mL) is transferred via peristaltic

pump. The same procedure is repeated for each of the remaining steps in the train: 10% inoculant

from the previous step and 90% fresh medium (of half of the bag volume).



In addition to the light energy mentioned above, each step of the train requires adequate

CO2. The CO2 will be provided using simple plastic tubing (much like those used at Genentech)

of 1” diameter. For the larger bags, 200L and above, multiple tubes will be placed to ensure

adequate CO2 transfer. Conveniently, the bags in question come equipped with filters to ensure

absolute sterility of any incoming substance. The CO2 requirement, as further discussed in the

previous section, “Medium Preparation”, is approximately 60% of the cell biomass. The

dissolved CO2 concentrations will be measured with a sensor. The pH of the trains should also be

measured and will be monitored and maintained at pH 6.8 with a standard probe. A sampling

probe will also be utilized to extract samples and monitor the cell density of each step of the

train. The bags have custom made probes to ensure absolute sterility and each probe

measurement will take place every 6 hours during each step of the train





5

Julio Baez.





14

Table 9: Innoculation Specifications

Innoculation Step 1 2 3 4 5 6

Wave Bag 20mL 200mL 2L 20L 200L 2000L

Dimensions

Innoculant Volume 1mL 10mL .1L 1L 10L 100L

Media Requirements 9mL 90mL .9L 9L 90L 900L

Light Source 34.46-51.7 W m-2 “ “ “ “ “

CO2 Supply

Time Required 16hrs “ “ “ “ “

Total Time for 96hrs

Innoculation

CO2 Supply Lines 1-1” silicon tube “ “ “ “ “



CO2 Valves membrane “ “ “ “ “

CO2 Pressure At pump-Positive displacement – “ “ “ “ “

Element release at 1000psi

CO2 Flow Element Magnetic meter : 0-62.3 L/min “ “ “ “ “

CO2 Probe Electrode: 10mbar minima “ “ “ “ “

pH Probe Electrode: 7.0+ 5 “ “ “ “ “

Light Probe Quantum:2-300 µmol m-2s-1 “ “ “ “ “

Sampling Probe Optical “ “ “ “ “

Transfer Pump Diaphragm: 0-62.3 L/min “ “ “ “ “

Transfer Tubes 1.5” silicon tubing “ “ “ “ “



Photobioreactor

In order to successfully produce 600kg of MAb per year, as the calculation highlighted in

the appendix illustrates, 343000 L of saturated cultivation solution will be required. This will

require 10- 10,000L tubular photobioreactors that are detailed below in Table 10. Every batch

will require two days: 1 to fill the tanks, and one to run the reaction. To meet the MAb demand,

we will perform 36 batches (fed batch) per year.



The tubular reactors currently used are in the order of 1cm-2.5cm (García-González et al.,

Mullikin and Rorrer). This is due to the light attenuation, modeled by Beer-Lambert Law,

exponentially decaying:



Iz = Io exp(-kXz)



Where: Iz PFD at z depth [μmol m-2s-]

Io PFD at surface [μmol m-2s-]

K Culture attenuation coefficient [L/mg cm]

X Concentration of biomass [mg dry weight/L]

Z Depth [cm]

Source: Tyler and Smith.







15

Because of this exponential reduction of light, small tubes are necessary. Calculations in

the appendix show that a 1 ½ inch tube is needed if light attenuation effects and pressure drops

are considered. While a 1.22 inch is best to reduce light attenuation, the pressure drop for a 2

hour filling cycle is 146645 Pa/m as compared to 46751.7 Pa/m in 1.5 inch pipes (see appendix for

all calculations).



The small diameter in the tube will require each reactor to have 15,860 m in length,

which will be a challenge to pump. If the photobioreactor was to be filled in 2 hours, the pressure

drop from pipe will be 8 million Pa –high pressure. For this reason, it is better if the

photobioreactor was filled over a 24 hour period to reduce the total pressure drop down to 0.5

million Pa. Using these sizing and flow rates, the total pressure drop will be 88 psi, which is

below the 1000 psi that diaphragm valves can generate.



In addition to satisfying pumping requirements, photobioreactor tubes should also be

resistant to the temperatures and pressures of the photobioreactor. Out of the many possibilities,

silicon flexible tubing was chosen because it is transparent, smooth, and it can support 88 psi.



Even though standard manufacturing practices using mammalian or bacterial cells require

the employment of impellers for mixing and jacket characteristics for adequate heat removal, our

design does not require the use of impellers or jackets. Algae in general prefer stagnant

environments and vigorous agitation hinders growth. Because light absorption is the only

limiting reagent in our reaction, and agitation does not increase light absorption, agitation is not

required for optimal growth. Dunaliella tertiolecta is also an alga that is capable to grow and

produce proteins optimally at ranges from 25-35ºC. Generally, algae thrive at higher

temperatures and therefore heat removal is an unnecessary expenditure. This allows our

organization to utilize a room thermostat from a standard HVAC to control temperature.



As mentioned previously, CO2 transfer will be accomplished using simple silicon tubing

of 1” diameter. Unlike CHO and E. coli expression systems, gas transfer is not a limiting reagent

in our algae process. The CO2 will be filtered upon entry into the photobioreactor to ensure

absolute sterility. In addition, our exit gas stream will also be filtered to eliminate outside

contamination.



The medium, detailed previously, will be fed into the photobioreactor in a fed batch

manner and filtered upon entry. The detailed process diagram is illustrated below in Figure 5.









16

Table 10: Tubular Photobioreactor Specifications

Tube Volume 10,000

Loading Volume 8,000

Wall Thickness 0.75 inches

Tube Material Flexible vitron tubing

Feed Type Fed Batch

Light Requirements 34.46-51.7 W m-2

CO2 Flow Rate 200 g/h

Media Flow Rate 6.9L/min

Time in tube 1 day (2 day if you count filling time)

Number of Cycles 36 per year

Productivity 35% of biomass

Innoculation Transfer Pipes 1.5” silicon tubing

Innoculation Transfer Pump (to Tube Diaphragm: 0-62.3 L/min

Photobioreactor)

CO2 Supply Lines 1” silicon tubing

CO2 Filters Hydrophobic, membrane-pleated: 0.01µm

rating

CO2 Control Valves Membrane valves, autoclavable

CO2 Pressure Element At pump-Positive displacement – release at

1000psi

CO2 Flow Element Magnetic meter : 0-62.3 L/min

CO2 Probe Electrode: 10mbar minima

Medium Supply Pipes Stainless steel, electropolished, 2.5”, tilted

Medium Filters Hydrophobic, membrane pleated: 0.02 µm

Medium Control Valves Membrane valves, autoclavable

Medium Flow Element Magnetic meter : 0-62.3 L/min

Sampling Probe Optical

pH Probe Electrode: 7.0+ 5

Light Probe Quantum:2-300 µmol m-2s-1

Tube Pressure Gauge At pump-Positive displacement – release at

1000psi









17

Figure 5: Photobioreactor P&ID



Deoxygenation

Oxygen removal is crucial for the photosynthetic activity of plant cells. Not only does

oxygen inhibit photosynthesis, but also high irradiance levels inside the photobioreactor with

high dissolved oxygen leads to photo-oxidation. A photobioreactor lacking dissolved oxygen

(DO) has an increased photosynthetic activity of 14%. Photosynthetic activity has been known to

decrease 35% when the DO content is 1.38 mol/m3 at 20ºC (Rubio et al. 1999). For these reasons

deoxygenating is required for photobioreactors. Currently, there are three ways in which this can

be achieved: Catatalytic oxygen removal systems (CORS), Activated Carbon Deoxygenation

(DEOX), and Gas Transfer Membranes (GTM).



CORS removes oxygen by treating media with hydrogen, and passing it trough a

palladium catalyst to produce water that is less than 1ppb in dissolved oxygen concentration.

DEOX, on the other hand, feeds a chemical oxygen scavenger known as hydrazine; the stream is

then directed to a ion-exchange resin vessel to removed trace impurities leached by the activated





18

carbon. This process yields a purity of 1ppb of dissolved oxygen. Finally GTM is the passing of

the stream through an impermeable membrane, but with gas transfer. This method also yields to

less than 1ppb in dissolved oxygen. Alternatively, hybrid methods have been suggested to reduce

oxygen to 7,(Kemp, 2004) and the contaminants are acidic using a cation

exchange column is appropriate. Here the desire proteins will bind to the column while the

contaminants and unwanted proteins will not be retain and leave as waste. Before every run the

column is sterilized with NaOH and regenerated with NaCl. Afterwards media from the gel

filtraion column is loaded on the column and allowed to bind. The proteins are then eluded with

a solution of .03M Na-Phosphate and .1M NaCl at pH 6.5. It is estimated that 95% of the

proteins will be recovered. In addition, samples of the out coming effluent will be tested every 6

hours with a sampling probe. The cation exchange column utilized is fully described in Table ,

and is illustrated in Figure 8. Additional equipment involved in this recovery step are described

in Table .



Anion Exchange Chromatography

Two anion exchange chromatography columns operated identically in parallel will

immediately follow cation exchange separation and are used for capturing the remaining trace

contaminants such as host cell protein and DNA residues from the protein media. Its main

function is to remove the negatively charged contaminants. It also separates all protein A leakage

products and removes small amount of oligomers from the antibody. Thus a large amount of host

cell proteins (hcp) will be bound while the neutral antibodies are allowed to flow through. The

typical contaminant clearance value for hcp removal is 10 for every ng/ml while for DNA it is

100 for every pg/ml (Kemp, 2004). At this step, the column is operated in the flow through mode

(Kemp pg. 93). Given that most antibodies are basic proteins, pI >7, (Kemp, 2004) and the

contaminants are acidic, using an anion exchange column is appropriate. In order to maintain

high volumetric flow rates and to prevent bottlenecking, the column is required to have a

minimum depth of approximately 10cm (Kemp, 2004, pg. 93). Determined by Necina et al.

(1993) the main ligand for the anion exchange column is the amino group. The anion exchange

column utilized is fully described in Table 7, and is illustrated in Figure 8. Additional equipment

involved in this recovery step are described in Table 8.

The anion exchange column is operated by first loading the incoming protein solution

from the protein A column onto the column. The contaminants will bind to the resin while the

protein will flow through. Before every run the column is regenerated with NaCl followed by a

NaOH wash. The mAb is then eluted using Na-Phosphate and NaCl solutions, pH 8.5. The

recovery rate is estimated to be 95%. A sampling probe will also be utilized to extract samples

every 6 hours during the anion exchange recovery step.









47

Table 7: Ion Exchange Operating Modules

CM- Hyper D Cation Exchange Mono Q Anion Exchange

Chromatography Chromatography (2)

Mass (kg) in per cycle 18.5 kg 8.77 kg

Mass (kg) out per cycle 17.5kg 8.3 kg

Recovery % .95 .95

Bed volume 2769.95 L 4154.92 L



Bed height 0.95m 1.09 m



Bed diameter 1.92 m 2.19m



Concentration of 6.67mg/mL 2.11mg/mL

protein in

Concentration of 2.11mg/mL 0.33 mg/mL

protein out

Conductivity Ranges 4-10 4-6

(AU)

Typical Buffer 8309 L of 0.03M Na-Phosphate, 12.5L of 50mM Na-Phosphate,

Media** 8309 L of 0.03M NaCl, pH 6.5 12.5L of 1M NaCl pH 6.3*

Buffer time 280 min 280 min

Elution Buffer 8309 L of 0.03 M Na-Phosphate 12.5L of 0.03 M Na-Phosphate

containing 8309 L of 1M NaCl, pH containing 12.5L of 1M NaCl,

6.5 pH 8.5

Elution flow rate 9.89 L/min 14.84L/min

Elution Time 280 min 280 min

Volumetric flow rate 0.0099 m^3 /min 0.014m^3/min

Resin type Fractogel EMD SE Hicap(M) beads Q-Sepharose, Preswollen in

20% ethanol

Resin pore size 800 A ~4,000,000 Da exclusion limit

Resin particle size 90 um 45 μm (wet)

Resin binding capacity 36mg of mAb mg/mL 30mg of mAb/mL

Resin storage NA 2-8°C

temperature

Particle size NA Not less than 95% than within

distribution 45 and 165 microns

Regeneration buffer 1-2 CV( or 2769L) of 1-2 M of 1-2 CV ( or 8309 L) of 1-

NaCL 2MNaCl followed by 1 CV ( or

8309 L ) of .1M NaOH in

8309L .5 MNaCL

Sanitization 2729L of .1-.5 M NaOH 8309L of.1-.5 M NaOH

Sanitization Flow rate 9.89L/min 14.84 L/min

Resin usage Needs to replace every 56 cycles Needs to be replaced every 20

cycles

Regenerate time 280 min 280 min







48

Resin Regeneration 1832.07mL/min 1044.23 mL/min

flow rate

Resin Volume 512978.4 mL 292397.7 mL

Linear gradient flow 9.892659 L/min 14.83899 L/min

rate

*Column outer 2.49m 2.85 m

diameter

*Column inner 1.91m 2.19m

diameter

*Column height 1.24 m 1.42 m

*Column Volume 3600.93 L 5401.40 L

Absorbance 280 nm 280nm

Number of cycles 36 cycles per year 36 cycles per year

Cycle time 14 hr 14 hr

pH condition range 1-13 2-12

Column Pressure drop 8 bar 1000 psi

limit

Actual estimated 5.47 bar 455.32 psi

Pressure drop of

column

Source: Sigma-Aldrich website, Subramanian,

http://www.merck.de/servlet/PB/show/1225330/w214154_Fracotgel_WF.pdf

*

taken to be 30% bigger than bed dimensions



Table 8. Additional Equipment for the Cation and Anion Exchange Column and their

Dimensions

Material Dimension Purpose

Pump 316L The pump removes solvents from the

elution, buffer, or regenerating tanks and

delivers it to the anion exchange column at

the appropriate flow rate

Na- 316L Diameter: 1.289m; Solvent used for the buffer and elution

Phosphate Height: 3.86m media for the cation and anion exchange

tank -5000L columns

NaOH tank- 316L Diameter: 1.289m; Solvent used for regenerating the anion

5000L Height: 3.86m column. Also use as a media for sanitizing

the cation exchange columns

NaCl tank- 316L Diameter: 1.289m; Solvent used for regenerating both the

5000L Height: 3.86m cation and anion column. Also used in the

buffer and elution media for the anion and

cation exchange columns









49

Elution 316L Diameter: 1.62m, It is also used for eluding the bound proteins

buffer tank Height: 3.25m from the cation exchange columns.

for cation-

10,000L

Elution 316L Diameter: 1.62m, Used for eluding the bound contaminants

buffer tank Height: 3.25m from the anion column. The contaminants

for anion- will leave the column and go to a waste

10,000L tank.

Buffer tank 316L Diameter: 1.62m, Use to maintain the appropriate pH

for anion Height: 3.25m

exchange

columns-

10,000L

Buffer tank 316L Diameter: 1.62m, Use to maintain the appropriate pH

for cation Height: 3.25m

exchange

columns-

10,000L

Regeneration 316L Diameter: 1.62m, Use to regenerate the column so that the

tank- Height: 3.25m contaminant from the incoming influent

10,000L stream will bind to the column

Air vent 316L To remove air from the highest point of a

coil or piping assembly

Valves 316L Use to monitor and control amount of fluid

leaving and entering

Sample Use to check the media in and leaving the

Probes anion and cation exchange column

UV and Monitor the absorbance (at 280nm) and

conductivity conductance of the product leaving the

monitor effluent stream.

Pressure Use to monitor the pressure of the influent

gauge

Air Vent For air bubbles or bypass

Piping Inner Diameter: 1.5 Transport media from one location to

1.5 inch inch another

Outer Diameter: 1.37

inch









50

Buffer Media .05M Na-Phosphate

(.03M Na-Phosphate,

Air .1M NaOH

.03 NaCL) .1M

Media from Vent

P-37

P-41 E-3

V-15 E-1

gel filtration P-42

P-29 P-35 NaCL

E-14 V-18

V-11 Elution Buffer P-20

P-43(.03M Na-Phosphate, P-30

E-6 Buffer Media

P-32

Pump V-23 1M NaCL,pH 6.5) P-46 PLC Elution Buffer ( 50mM of Na-Phosphate,

P-22

Pressure S-4 S-3

E-10

(.03M Na-Phosphate, .1MNaCL)

Gauge P-45

P-19

I-1

1MNaCL, pH 8.5) pH probe

P V-6 S-1 PLC

P-12

I-8 pH probe S-6



E-15 S-2

V-14

S-5 V-24

pH probe E-13

P-5

I-2



I-16 V-3



Influent PLC

P-23 E-11

V-13 V-2

E-9 P-3 pH probe

pH probe P-8

P-18



pH probe V-1 V-10

P-28

Regenerant NaCL



P-47 P-55

pH probe P-1

P-34 P-31



Computer LabView Software

P-43 P-2







Anion Exchange P-15 P-11

P-40 pH probe P-17

Column

I-6

V-25 V-7

Cation exchange P-44

Column

Anion Exchange

Column

E-7

E-5

P-38







Sampling Probe UV Monitor

Target

P-36

Protein

P-57 I-4







V-22 Conductivity Monitor

P-27

P-54



V-5

V-20

E-12 E-2

waste T V-9

E-16 V-16

Target I-3 P-33

Conductivity Cell V-4

Protein

Waste P-27 Target P-6 P-21 P-7 P-14

P-49

I-10

Protein Conductivity Cell

P-24

Sampling Probe V-19 T

V-17 I-13

Waste UVI-7Cell

V-21 I-5

Conductivity Monitor Sampling Probe Target Protein

P-27 P-10

T



P-50 T

I-15

UV Cell

I-11

P-52









Computer LabView Software

I-12

UV Monitor

I-14



P-51









E-4









Figure 8. P&ID of Cation and Anion Exchange Columns









51

Downstream CIP/SIP Systems__________________________



For the downstream process, the same multi-use cleaning system that was used to clean

the upstream units will be utilized. The figure of the multi-use cleaning system is available in

part 2 of the project. Again cleaning reagents are dispersed through the circuits for different

intervals of time at varying temperatures. The specifications for the CIP/SIP systems are

described in Table 29. A typical cleaning schedule for our system is illustrated in Table 30.



The water for injection is used to supply the water necessary in the CIP and SIP systems.

Combined, these processes required 54,520 L of WFI. Before using the water, it first needs to be

sterilized. The water is carbon filtered and then undergoes reverse osmosis. The reverse osmosis

membrane captures ions, microbes, endotoxins, and other small contaminants (Lyderson, pg.

548).









52

Table 29: CIP/SIP Specifications for Downstream Processing



Material Description Dimension

Scavenge Return Pump The return pump pumps the fluids out of the

clean process equipment, and either returns

it to the CIP unit (recycling) or pumps it

directly to the drainage system.

Double Pipe Heat exchanger The built-in heating elements heat the fluid

to a specified temperature.

Piping 316L SS Use to connect equipment to one another 1.5 inch

Sodium hydrochlorite 10000L 316L SS Detergent for CIP. Use to id media and Diameter:

tank algae residues 1.62m, Height:

4.86m

Water tank 10000L 316L SS For multiple cleaning and solution Diameter:

preparation purposes. 1.62m, Height:

4.86m

Sodium hydroxide 10000L 316L SS Chemical for cell disruptor. Diameter:

1.62m, Height:

4.86m

Sodium hydrochloride 10000L 316L SS Chemical for cell disruptor. Diameter:

1.62m, Height:

4.86m

Acetone 100L 316L SS Chemical for centrifuge. Diameter: 0.34,

Height: 1.05



H3PO4 10000L 316L SS Protein A Column Diameter:

1.62m, Height:

4.86m

Guanidine hydrochloride10000L 316L SS Protein A Column Diameter:

1.62m, Height:

4.86m

C2H6O10000L 316L SS Ion Exchange Column Diameter:

1.62m, Height:

4.86m

0.05 M CH2O2 in 40% C2H6O and 316L SS Ion Exchange Column Diameter:

60% H2010000L 1.62m, Height:

4.86m









53

0.2 NaH2PO4 + 0.3 M 316L SS Ion Exchange Column Diameter:

CH3COONa10000L 1.62m, Height:

4.86m

Water 10000L tank 316L SS Water comes from WFI . Use for cleaning with Diameter:

and for buffer media 1.62m, Height:

4.86m

Sodium hydroxide 10,000L tank 316L SS Second solvent used for cleaning purposes with Diameter:

1.62m, Height:

4.86m

Control valves Use to monitor and control amount of fluid

leaving and entering

Cold Water Supply tank (3000L) 316L SS Supply the system with cool to cold water with Diameter:

1.08m, Height:

3.25m

Steam Supply tank (3000L) 316L SS Supply heat exchanger with steam with Diameter:

1.08m, Height:

3.25m

Condensate Return box (200L) 316L SS Return the collected condensate amount to with Diameter:

heat exchanger .4m, Height:

1.3m

Pressure relief valve Use to protect system’s pressure vessels

Duplex Filter Filter steam from heat exchanger

PLC (include software and

printer) Built-in PLC unit with an operating panel.

Placed in the control room.

The PLC is programmed individually for

each cleaning task. Each task is fully

documented by a laser printer.

Sample probe Use to check the detergent and buffer media

Positive Displacement Pump Use for injecting chemicals through an the

pipeline system at precise rates.

CIP pump The frequency controlled supply pump

builds up the pressure to the spray nozzle

and washing spear, that is fitted in the

process equipment to be cleaned.

Supply Vessels 316L SS Holds the cleaning reagent and or water for

the unit to be clean

Compressed Air Filtered compressed air purge system

ejects residual water from the lines.

WFI Carbon Filter Water scrub unit, Model: WSU1000P 60gpm capacity



WFI Reverse Osmosis Membrane Osmonics, Model: M-CB4040AD 4gpm capacity

1.0µm rating

Sources: ACCompacting, MorkUSA, LHS Company

(Note: Assume 3:1 Height to Diameter ratio for Tank dimensions)







54

Table 30: Typical Cleaning Schedule

(Note: An air blow step occurs in between each of the cleaning step)

Equipment Operation Flow Amount Time Temperature Pressure

rate

Cell Water Rinse 38 L/min 760 L 15-20 Ambient 2-8 bar

disruptor min (water

feeding)

Sodium 38 L/min 15.048 of 15-20 Ambient to

hydroxide wash NaOH + min 75 ºC

(0.1, ) 135.4 gal of

water

Water Rinse 38 L/min 570L or 5-15 Ambient 2-8 bar

150.48 gal min (water

feeding

pressure)

Sodium 38 L/min 15.048 gal of 15-20 Ambient to

hydrochloride Sodium min 75 ºC

wash Hydrochlorite

+ 135.43 gal

of water

Water Rinse 38 L/min 380 L or 5-10 Ambient 2-8 bar

100.32 gal min (water

feeding

pressure)

Purged/Drainage 38 L/min ----- 10-15 ------

min

Steam sterilized 130 lb/hr 43 lb 5-20 Ambient to 3-5 bar

min 100C

Centrifuge Water rinse 10 L/min 10L 1 min 100 ºC 2-8 bar

(first step) (water

feeding)

Water flood 10 L/min 10L 1 min 100 ºC 2-8 bar

(water

feeding)

Water rinse 10 L/min 6L 0.6 min Ambient 2-8 bar

(water

feeding)

Acetone flood 10 L/min 5L 0.5 min Ambient 2-8 bar

(water

feeding)

Water rinse 10 L/min 40L 4 min 100 ºC 2-8 bar

(water

feeding)









55

Centrifuge Water rinse 10 L/min 40L 4 min 100 ºC 2-8 bar

(2nd step) (water

feeding)

Water flood 10 L/min 40L 4 min 100 ºC 2-8 bar

(water

feeding)

Water rinse 10 L/min 30L 3 min Ambient 2-8 bar

(water

feeding)

Acetone flood 10 L/min 25L 2.5 min Ambient 2-8 bar

(water

feeding)





Water rinse 10 L/min 40L 4 min 100 ºC 2-8 bar

(water

feeding)





All 1000L Water Rinse 38 L/min 760 L 15-20 Ambient 2-8 bar

tanks-ie min (water

Holding feeding)

Tank,

Elution tank,

Regeneration

tank, Buffer

tank, Citrate

tank, Tris-

HCL tank

Sodium 38 L/min 15.048 of 15-20 Ambient to

hydroxide wash NaOH + min 75 ºC

(0.1, ) 135.4 gal of

water

Water Rinse 38 L/min 570L or 5-15 Ambient 2-8 bar

150.48 gal min (water

feeding

pressure)

Sodium 38 L/min 15.048 gal of 15-20 Ambient to

hydrochloride Sodium min 75 ºC

wash Hydrochlorite

+ 135.43 gal

of water









56

Water Rinse 38 L/min 380 L or 5-10 Ambient 2-8 bar

100.32 gal min (water

feeding

pressure)

Purged/Drainage 38 L/min ----- 10-15 ------

min







All 5000L Water Rinse 38 L/min 760 L 15-20 Ambient 2-8 bar

tank- NaCl min (water

tank, NaOH feeding)

tank, Na-P

tank

Sodium 38 L/min 15.048 of 15-20 Ambient to

hydroxide wash NaOH + min 75 ºC

(0.1, ) 135.4 gal of

water

Water Rinse 38 L/min 570L or 5-15 Ambient 2-8 bar

150.48 gal min (water

feeding

pressure)

Sodium 38 L/min 15.048 gal of 15-20 Ambient to

hydrochloride Sodium min 75 ºC

wash Hydrochlorite

+ 135.43 of

water

Water Rinse 38 L/min 380 L or 5-10 Ambient 2-8 bar

100.32 gal min (water

feeding

pressure)

Purged/Drainage 38 L/min ----- 10-15 ------

min







2000L Water Rinse 38 L/min 760 L 15-20 Ambient 2-8 bar

Affinity Prep min (water

Protein A feeding)

container









57

Sodium 38 L/min 15.048 of 15-20 Ambient to

hydroxide wash NaOH + min 75 ºC

(0.1, ) 135.4 gal of

water

Water Rinse 38 L/min 570L or 5-15 Ambient 2-8 bar

150.48 gal min (water

feeding

pressure)

Sodium 38 L/min 15.048 gal of 15-20 Ambient to

hydrochloride Sodium min 75 ºC

wash Hydrochlorite

+ 135.43 of

water

Water Rinse 38 L/min 380 L or 5-10 Ambient 2-8 bar

100.32 gal min (water

feeding

pressure)

Purged/Drainage 38 L/min ----- 10-15 ------

min







Protein A Acid wash- 64 L/min 7700L 15-20 Ambient 2-8 bar

Column H3PO4, pH 1.5 min (water

feeding)

Water flood-4-6 22 L/min 7700L 15-20 Ambient 2-8 bar

M Guanidine min (water

hydrochloride feeding)

Storage - C2H6O 22 L/min 3080L 15-20 Ambient 2-8 bar

min (water

feeding)

Ion Cleaning-0.05 22 L/min 21,000L 15-20 Ambient 2-8 bar

Exchange M CH2O2 in min (water

Column 40% C2H6O and feeding)

60% H20

Water wash- 22 L/min 21,000L 5-10 Ambient 2-8 bar

H20 min (water

feeding)

Conditioning- 22 L/min 21,000L 15-20 Ambient 2-8 bar

0.2 NaH2PO4 + min (water

0.3 M feeding)

CH3COONa









58

Salt removal- 22 L/min 21,000L 5-10 Ambient 2-8 bar

H20 min (water

feeding)

Non-specific 22 L/min 1400L 15-20 Ambient 2-8 bar

saturation- min (water

Gelatin feeding)









Cost Analysis of Downstream Process Design





The five sections of the downstream processing include: centrifugation, cell disruption,

another centrifugation step, two protein A affinity columns operated in parallel, a gel filtration

column, a cation exchange column, and two anion exchange chromatography columns operated

in parallel. The cost associated with both installation, maintenance, and operation of each section

have been determined and listed in the following tables. The final yearly operational cost of the

downstream side of the process sums up to be $ 14.1 Million, while the one time start up cost is

estimated to be $10.3 Million. This cost does not take into account employees compensation or

building costs. This cost also does not yet include the cost of the gel filtration column.



Cell Harvest : Centrifugation



Table 31: Fixed Capitol Costs for Cell Harvest Centrifugation

Item Quantity Unit Total Cost

Cost

BRPX617SFV, 20L, 8000rpm, 60cm Diameter 2 $500,500 $1,001,031

Diaphragm Valve 2 $449 $898

Butterfly Valves (CIP) 6 $300 $1,800

Centrifugal Pump 8 $3,000 $24,000

Process Piping (20% of major equipment cost, Atkinson __ __ $200,206

1991)

Installation (25% of major equipment, Atkinson, 1991) __ __ $250,258



Validation (20% of major equipment, Atkinson, 1991) __ __ $200,206



Shipping (40% of major equipment cost, Atkinson, 1991) __ __ $400,000



Total Capitol Investment for Cell Harvest Centrifugation $2.1 million









59

Table 32:Yearly Operational Cost for Cell Harvest Centrifugation

Item Quantity per Year Unit Cost Total

Cost

Maintenance 2 3% of Capitol Investment $45,000

(Atkinson, 1991)

Cooling Water 100,000L, 36 times per year: $0.50 per Liter $1,800

3,600,000L

Biological Waste 66% of 100,000L, 36 times per $0.20/100L $4783

Treatment year: 2.4million Litres

Steam Sterilization 500tonne $15/tonne $7500

Energy Consumption 35,840 kWh $0.09/kWh $3226

Total Yearly $62,310

Operational Costs





Cell disruption



Table 33: Capitol Costs for Bead Mill Operation

Item Unit Quantity Total Cost

Cost

316L s.s. Dyno-Mill MULTI-PILOT $229,000 5 $1.15

305 Liter , 5000rpm, 70L/min million

316L Buffer Solution Holding Tank, 381 Liter $900 1 $900

316L Mixing Tank, 1906 Liter $4000 1 $4000

Installation Costs (25% major equipment cost, __ __ $286,000

Atkinson, 1991))

Process Piping (20% of major equipment cost, __ __ $230,000

Atkinson, 1991)

Splitters $100 7 $700

Diaphragm Valve $449 4 $1,800

Centrifugal Pump $3,000 7 $31,000

Shipping (40% of major equipment, Atkinson, 1991) __ __ $460,000



Validation (up to 20% of major equipment, __ __ $230,000

Atkinson, 1991)

Total Fixed Capitol Costs for Bead Mill $2.4 million









60

Table 34: Yearly Operational Costs for Bead Mill Operation

Item Unit Cost Quantity per Year Total

Cost

Maintenance 3% of Capitol 2 $72,000

Investment, Atkinson,

1991

Cooling Water $0.50 per Liter $34,000L, 36 times per $612,000

year: 1.2 million liters

Buffer Solution (plus $3.00 per liter $8,400 L. 36 times per $907,200

reducing agents and year: 302,400 liters

inhibitors)

Energy Consumption $0.09/kWh 40,000 kWh $3,600

Total Yearly Operational $1.6 million

Costs for Bead Mill





If the lack of cell wall on our algae strain results in easy cell rupture, an alternative , more

cost effective Microfluidizer system can be used instead. Homogenizers were priced around

$278,000 in 1986. This equipment, adjusted for 3% annual inflation, will cost approximately

$500,000 in 2006. Because the microfluidizer system includes the control elements,

programmable logic controllers, and heat exchanger, the cost analysis for cell disruption will not

require the addition of these elements. The only additional cost associated with this system is the

cost of electricity and the cost for running the heat exchanger, which was estimated to annually

cost $19000 and $4000, respectively. For calculating total capitol investment as well as selling

price required, we will use the more costly bead mill system to ensure proper budgeting.



Cell Debris Removal: Centrifugation

Table 35: One Time Costs for Cell Debris Removal Centrifugation

Item Quantity Unit Cost Total Cost

BRPX617SFV, 57L, 9,000rpm, 100cm Diameter 2 $1,430,044 $2,860,088

Diaphragm Valve 2 $449 $898

Butterfly Valve (CIP) 6 $300 $1,800

Centrifugal Pump 8 $3,000 $24,000

Process Piping (20% of major equipment, Atkinson, 1991) __ __ $572,018



Installation (25% of major equipment , Atkinson, 1991) __ __ $715,022



Shipping (40% of major equipment, Atkinson, 1991) __ __ $1.2

million



Valladation (20% of major equipment, Atkinson, 1991) __ __ $580,000



Total Fixed Capitol Costs for Cell Debris Removal $6.0 million

Centrifugation





61

Table 36 :Yearly Operational Cost for Cell Debris Removal Centrifugation

Item Quantity per Year Unit Cost Total

Cost

Maintenance 2 3% of Capitol Investment $126,000

(Atkinson, 1991)

Cooling Water 100,000L, 36 times per year: $0.50 per Liter $1,800

3,600,000L

Biological Waste 5% of 100,000L, 36 times per $0.20/100L $360

Treatment year: 180,000 Litres

Steam Sterilization 500tonne $15/tonne $7500

Energy Consumption 35,840 kWh $0.09/kWh $3226

Total Yearly $138,886

Operational Costs





Affinity Column

Table 37: Cost of Equipment and Operation of Affinity Column.

Equipment Units Needed Cost per unit

Affinity Column Shell 2 Columns $8,032.55*

Holding Tank 1 $20,855.36*

Resin 1540 L per year $6,000 per liter of affi-prep protein A**

Tris-HCl 180,000 L per $1.47 per liter***

year

Sodium Citrate 180,000 L per $0.08 per liter****

year

Adsorption spectrophotometer 2 $5,000

equipped with LabView to

monitor adsorption

Centrifugal Pump 1 $707.28

Motor for Pump 1 $1850

Pressure Meter 2 $500 per meter

Temperature probe 2 $260 per probe

pH Probe 2 $550 per probe

Flow Meter 2 $125 per meter

Tris-HCl Buffer Container 1 $20,855.36*

Sodium Citrate Elutant 1 $20,855.36*

Container

Affi-prep Protein A Container 1 $9,177.89*

Diaphragm Valve 12 $449.30 per valve

Diaphragm Membranes 48 per year $100 per membrane

Sampling Probe 2 $1095.00 per probe





62

Piping (20% of major $32,000

equipment)

Maintenance (3% of capital $5,000

costs)

Installation (25% of major $40,000

equipment)

Total Fixed Capital Costs $162,320.00

Total Annual Operational $9.5 Million

Cost

*

(Seider, 2004)

**

(Harrison, pg. 367)

***

(Bio-Rad website)

****

(Sigma-Aldrich website)





Cation Exchange Columns



Table 14. Cost of Cation Exchange Column

Equipment Type Quantity Cost (for each) Total cost

Resin (replace Fractogel EMD 512978.4 mL $457053.5

every 56 cycles) SE Hicap(M)

beads $445.49 for

500 mL

Column Cation 1 $374,000 $374,000

Membranes for 28 $100 each $2800

valves

Elution Buffer 10,000L; 316L 1 $20,000 $20,000

tank

Buffer tank 10,000L; 316L 1 $20,000 $20,000

Computer 1 $4095.0 $8190

Program

Software (ie Lab

view)

UV and PC 1 $2105 $4210

Conductivity

monitor

Sample probes PFA 100 1 $1095 $6579

Pumps Diaphragm 1 $5000 each $5000



Valves High Purity Two 7 $449.30 $3143

Way diaphragm

Valve- EV-98613-

56









63

Pressure gauge High-accuracy 1 $800.0 $800.0

digital test gauge,

0 to 30" Hg range

EW-68330-00

Electricity 856 kW* .8 cent per kW



$68.48

per cycle

316L Piping (1.5 Estimated to be $ 18000

in) 20% of major

equipment

NaOH cost 5538 L $63.78 per 18L $ 2211.04 per

cycle

NaCl cost 19387 L $81.70 per 300mL $45316.67 per

cycle

Na-Phosphate 16618 L $51.0 for 2.5kg $ 16972.8 per

cycle

Motor for Pump 32" Commerical $1849.95 $1859.95

Mower with 10.5

B&S Engine 1

Pressure Release At pump 6- come with $0 $0

pump

Conductivity Cell 1 $5000 $5000

UV Cell 1 $5000 $5000

Direct Soft PLC 1 $395.0 each $790

software

PLC controller 2 $1095.0 $2190

STV Sanders 1 $1573.0 $3146

Pressure relief

valve

Maintenance Estimated to be $2700 per year

3% of major

equipment

Installation Estimated to be $22500

25% of major

equipment

Shipping Estimated to 40% $36000

of major

equipment

Validation Estimated to be $18000

20% of major

equipment

Total Fixed .54 Million

Capital Costs









64

Total 7.5 Million

operational cost





Anion Exchange Columns

Table 14. Cost of Anion Exchange Columns

Equipment Type Quantity Cost (for each) Total cost

Resin (replace Q-Sepharose, 555666.67 mL $852392.66

every 20 cycles) Preswollen in per column

20% ethanol $383.50 for

500mL

Column Anion 2 $254,000 $508,000

Membranes for 72 $100 each $7200

valves

Regeneration 10,000L; 316L 1 $20,000 $20,000

Buffer tank

Elution Buffer 10,000L; 316L 1 $20,000 $20,000

tank

Buffer tank 10,000L; 316L 1 $20,000 $20,000

Computer 2 $4095.0 $8190

Program

Software (ie Lab

view)

UV and PC 2 $2105 $4210

Conductivity

monitor

Sample probes PFA 100 6 $1095 $6579

Pumps Diaphragm 5 $5000 each $25,000



Valves High Purity Two 18 $449.30 $8087.4

Way diaphragm

Valve- EV-

98613-56

Na-Phosphate 5,000L; 316L 1 $10,000 $10,000

tank

NaCl tank 5,000L; 316L 1 $10,000 $10,000

NaOH tank 5,000L; 316L 1 $10,000 $10,000

Air Vent 1 $1200 $1200

Pressure gauge High-accuracy 1 $800.0 $800.0

digital test gauge,

0 to 30" Hg range

EW-68330-00









65

Electricity 856 kW* .8 cent per kW



$68.48

per cycle

316L Piping (1.5 300ft Estimated to be $ 124600

in) 20% of major

equipment

NaOH cost 16618 L per $63.78 per 18L $ 58883 per

cycle cycle

NaCl cost 16643 L per $81.70 per $4532443 per

cycle 300mL cycle

Na-Phosphate 832L per cycle $51.0 for 2.5 L $ 16972.8 per

cycle

Motor for Pump 32" Commerical $1849.95 $9249.75

Mower with 10.5

B&S Engine 5

Pressure Release At pump 6- come with $0 $0

pump

Conductivity Cell 1 $5000 $5000

UV Cell 1 $5000 $5000

Direct Soft PLC 2 $395.0 each $790

software

PLC controller 2 $1095.0 $2190

STV Sanders 2 $1573.0 $3146

Pressure relief

valve

Maintenance Estimated to be $18690 per

3% of major year

equipment

Installation Estimated to be $155750

25% of major

equipment

Shipping Estimated to be $249200

40% of major

equipment

Validation Estimated to be $ 124600

20% of major

equipment

Final Fixed $2.9 Million

Capital Costs

Total $5.3 Million

Operational

Cost

Source: Sigma Aldrich website, Harrison

*

The electricity cost is estimated as 20% of the process equipment cost







66

**

The maintenance and installation cost is estimated as 3% of the process equipment cost









CIP/SIP for Downstream Process



Table 39: Cost of CIP/SIP system for downstream process

CIP materials Quantity Cost per unit Total Cost

316L piping 1500ft $14.07 ft $21,105

Electricity 515kW $ 203/kW $104545

NaOH 58.75 L per cycle $5.16 per kg $30315

H20 54,519.37 L per cycle $.50 gallon $2710

H3PO4, pH 1.5 10000 L $466 per gallon $1211600

CH3COCH3 30L per cycle $87.4 per gallon $2644

CH6ClN3 3840 kg per cycle $354.3 per kg $1360512

C2H6O $16000 L per cycle $40.9 per gallon $170144

CH2O2 69kg per cycle $151 per gallon $10419

NaH2PO4 983 kg per cycle $172 per kg $169144.8

CH3COONa 738 kg per cycle $125 per kg $92250

Gelatin 3kg per cycle $138 per kg $615

Total Fixed Capital $21,105

Costs

Total Operational Cost $ 3.1 Million









67

Economic Feasibility of MAb Production

in

Dunaliela tertiolecta







Fixed and Operational Costs of Entire Process____________



The total process for producing monoclonal antibodies includes both an upstream and a

downstream. The upstream includes four steps: medium preparation, inoculation train,

photobioreactor, and CIP/SIP. The downstream includes seven steps: centrifugation, cell

disruption, another centrifugation step, two protein A affinity columns operated in parallel, a gel

filtration chromatograph column, a cation exchange column, and two anion exchange

chromatography columns operated in parallel. The total cost of the process is summarized in

Tables 40 and 41 below. The total one time cost for the entire process is $13.8 million and the

total annual operational cost is $18.9 million.



Table 40: Total One Time Costs for Entire Process

One Time Costs Cost ($)

Upstream Process $3.5 Million

Downstream Process $10.3 Million

Total Fixed Capital Costs for Entire Process $13.8 Million



Table 41: Total Annual Operational Costs for Entire Process

Annual Operational Costs Cost ($)

Upstream Process $6.1 Million

Downstream Process $14.1 Million

Total Annual Operational Costs for Entire $20.2 Million

Process







68

Selling Price to Receive 20% Return on Investment________





The following tables show potential sell prices to receive the minimum 20% return on

investment (ROI) for the first year of production (Table 42) as well as years after (Table 43). We

have included 50% contingency and 10% interest rates on money borrowed to account for the

high risk business venture we are pursuing. We have also accounted for employee salaries,

pensions, and healthcare. Considering that current contract manufacturing procedures result in

MAb selling prices between $2000 and $5000, our proposed algal process is a very competitive

business venture.



Table 42: Sell price per gram of MAb required to receive 20% ROI for first year of

production

Total Annual Operating Costs (includes maintenance and utilities)

Total Capitol Investment (includes shipping and validation)

Annual Labor and Payroll (300 employess with avg. salary of $100,000 per year to $30

include pensions and healthcare) million

Start-up Expenses (10% of capitol investment, Atkinnson, 1991)

Contingency (50% of capitol investment for unventured process, Atkinson, 1991)

Other Plant Overhead (telephone, sewer, freight, travel, legal, rentals) $400,000

Taxes and Insurance (6% of capitol investment, Atkinson 1991)

Annual Expense due to Interest Rate for Capitol Investment (10%)

Sales and Marketing

Total costs for first year of operation

Price per gram MAb to receive minimum 20% ROI



Table 43: Sell price per gram of MAb required to receive 20% ROI every other year of

production



Total Annual Operating Costs (includes utilities) $20.2

million

Annual Labor and Payroll (300 employees with avg. salary $100,000 per year to $30 million

include pensions and healthcare)

Contingency (50% of capitol investment for un-ventured process, Atkinson, 1991) $6.9 million

Other Plant Overhead (telephone, sewer, freight, travel, legal, rentals) $400,000

Taxes and Insurance (6% of capitol investment, Atkinson 1991) $828,000

Annual Expense due to Interest Rate for Capitol Investment (10%)

Sales and Marketing





69

Total annual costs after first year $58.3

million

Price per gram MAb to receive minimum 20% ROI $117









Nomenclature





Symbol Meaning

F Force

Q Charge

E Dielectric constant

r Radius

Vg Sedimentation

velocity

n Plate number

w Angular velocity

g Gravity constant

Σ Sigma Factor

a Area

ρ Density

μ Viscosity









70

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73

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Appendix

Calculations for Process Capacity, Bioreactor Design, Media Requirements, and Flow

Rates will be in the Final Report.



Upstream Media Sterilization Heat Exchanger Calculations

assume a volume of 10,000 Liters of media needs to be sterilized

sterilization of 10 m3

Time for sterilization 2 hr

Flow rate 1.389 L/s 0.001389 m3/s

KJ/(Kg

specific heat of water 4.186 *K)

KJ/(Kg

specific heat of sea water 3.93 *K)





Regen. Heat Exchanger



T hot in 130 C

T hot out 50 C

T cold in 30 C

T cold out 120 C



choose plate data

width, w 40 cm 0.4 m

height, h 30 cm 0.3 m

spacing 0.4 cm 0.004 m

number of

plates 16 plates

port size 30 cm

Thickness of the plates, a 0.00117 m





delta T ln 358.0519

delta T 373.15 K

Q 2169.453 KW

Q delta T ln 2081.674









74

A 1.68 m2

U 3460.648 W/(m2*K)





thermal conductivity

k (hot stream) 0.016 W/(m*K)

k(cold stream) 0.58 W/(m*K)

k (SS) 17 W/(m*K)



Pr=Cp*u/k

Pr (hot stream) 196.5

Pr (cold

stream) 6.775862



Re 2320.719



Nu=0.023*Re^(.8)*Pr^(1/3)

Nu(hot stream) 718.4363

Nu(cold

stream) 233.8419





hydraulic diameter

Dh 0.002333



h=Nu*k/Dh

h(hot) 4926.754 W/(m2*K)

h(cold) 58130.34 W/(m2*K)





1/U 0.000289 (m2*K)/W



U 3460.212 W/(m2*K)

assume diameter of pipe is 3 inches 0.0762 m

cross sectional area of

pipe 0.00456 m2

velocity = 0.304556 m/s









Steam Heat Exchanger



T hot in 140 C

T hot out 120 C

T cold in 120 C

T cold out 130 C



choose plate data

width, w 40 cm 0.4 m

height, h 30 cm 0.3 m







75

spacing 0.4 cm 0.004 m

number of

plates 24 plates

port size 30 cm

Thickness of the plates, a 0.00117 m





delta T ln 287.577

delta T 293.15 K

Q 1704.342 KW

Q delta T ln









A 2.64 m2

U 2202.231 W/(m2*K)





thermal conductivity

k (hot stream) 0.016 W/(m*K)

k(cold stream) 0.016 W/(m*K)

k (SS) 17 W/(m*K)





Pr (hot stream) 209.3

Pr (cold

stream) 245.625



Re 2320.719





Nu(hot stream) 733.709

Nu(cold

stream) 773.9121





hydraulic diameter

Dh 0.002333





h(hot) 5031.488 W/(m2*K)

h(cold) 5307.185 W/(m2*K)





1/U 0.000456 (m2*K)/W



U 2193.003 W/(m2*K)

how much steam is needed







76

length of holding tube = v* 5 mins Q= 1704.342 kW

0.0762

diameter of holding tube 3 inches m m dot = 5.813889 L/s

cross sectional area of

tube 0.00456 m2 for .5 hrs,

Liters of

steam per

heat

velocity in sterilized

tube 0.304556 m/s m= 10465 batch

length of tube 91.36691 m

Upstream CIP Calculations

CIP Appendix



Velocity= ,408 where Q is the flow rate and d is the

Reynolds number = (d*v*p)/mu Q/d^2 diameter



d inner diameter 2 inch

v fluid velocity 6.51 ft/s or 122.9609 m/min

p fluid density 60.6 lb/ft^3

flow

mu viscosity 0.35 cP Line size d inside Re rate

30 L/min or

1.5 1.37 243013 38gpm





For

steam

30 gpm or 14434.62 lb/hr Re 192000 for inner diameter 1.37





Cost Analysis water(L) sodium hydroxide

1liter= .264 gal 760 49.6 gal

684 Sodium hydrochlorite

570 49.6 gal

steam

380 58400 lb

684

380



380

380





380

total water 4598 L or 1213.872 gal of water









flow rate of sodium hydrochlorite and sodium hydroxide

Re= 6.31 *W/ d * mu where W is lb/hr









77

Re 4000

W 303.962

flow rate 0.625481 gpm







Centrifugation Calculations: Cell Removal



First Centrifuge: Remove Cells from Broth





Cell Density Calculations



Vcell

cell D (um) dry cell (g) water/cell mass (g) (um3) Vcell(cm3) cell density (g/cm3)=p

1.10E+0

10 1.15E-10 0.8 5.75E-10 523.3333 5.23E-10 0





Disc-Type Centrifugation Equation: Maximum Feed Rate





Q



 2a 2    o g  2n 2 Ro 3  R13 cot  





  v g 

 9 

 3g 

 bowl radius

theta (cm)

42 30



u Ro Ri

a (cm) p(g/cm3) po(g/cm3) g (cm/s2) g/(cm*s) n w(rad/s) (cm) (cm)

0.0005 1.10E+00 1 980 0.798 100 628 25 8



Q (l/min) Q (m3/min) Feed Rate From Fermentation

70 0.07



Q (cm3/s) vg sigma

3742.539 6.7357E-06 5.56E+08

0.224552 Maximum Feed Rate Allowed to Achieve Seperation (m3/min)



Therefore, our process flow rate is acceptable to achieve desired separation in the chosen

centrifuge.









78

Time Required to Achieve Seperation



 2 ro

Gt  t Generalized Gt for Algae Cells

g 3.00E+05



w(rad/s) Ro g t(seconds)

29.81865

628 25 980 39





Centrifuge Capacity Centrifuge Height



Volume(L) Q(L/min) time to separate (s) Volume Cone = 1/3*pi*r2*H (m3) r(m) h(m) h(cm)

34.78843 70 29.81865 0.034788 0.25 0.53 53.18



Centrifuge Time For Total Process Batch Volume of 100,000L



Batch Volume per

Q(L/min) Reactor Total Batch Volume Total time (min) Time (hours)

23.80

70 10,000 100000 1428.571 952





Second Centrifuge: Remove Cell Debris from Protein Broth

Cell Density Calculations



Vcell

cell D (um) dry cell (g) water/cell mass (g) (um3) Vcell(cm3) cell density (g/cm3)=p

10 1.15E-10 0.8 5.75E-10 523.3333 5.23333E-10 1.10E+00

debris mass debris V cell debris density

7.48E-11 6.54167E-11 1.14E+00

Disc-Type Centrifugation Equation: Maximum Feed Rate





Q



 2a 2    o g  2n 2 Ro 3  R13 cot  





  v g 

 9 

 3g 

 theta bowl radius (cm)

42 40









79

u

a (cm) p(g/cm3) po(g/cm3) g (cm/s2) g/(cm*s) n w(rad/s) Ro (cm) Ri (cm)

0.00025 1.14E+00 1.07 980 0.798 100 837.3333 35 8



Q (l/min) Q (m3/min) Feed Rate From Fermentation

70 0.07



Q (cm3/s) vg sigma

3432.21 1.24E-06 2.769E+09



0.205933 Maximum Feed Rate Allowed to Achieve Separation (m3/min)

Therefore, our process flow rate is acceptable to achieve desired separation in the chosen

centrifuge.

The maximum throughput of the centrifuge will allow for fluctuations of up to 75% more of the expected process flow rate.

Time Required to Achieve Seperation



 2 ro

Gt  t Generalized Gt for Algae Cells

g 2.00E+06



w(rad/s) Ro g t(seconds)

837.3333 35 980 79.87139





Centrifuge Capacity Centrifuge Height



Volume(L) Q(L/min) time to separate (s) Volume Cone = 1/3*pi*r2*H (m3) r(m) h(m) h(cm)

93.18329 70 79.871394 0.093183 0.35 0.726764 72.67643



Centrifuge Time For Total Process Batch Volume of 100,000L



Batch Volume per

Q(L/min) Reactor Total Batch Volume Total time (min) Time (hours)

70 10,000 100000 1428.571429 23.80952









Lysis by microfluidizer



ln(1-R) = -kNbPa



b = 0.3

k = 0.3 MPa-a

a = 0.6

p = 20,000 psi = 137 MPa

R = Extent of disruption

N = Number of pass



Ln(1-R) = -5.74

1-R = exp (-5.74)





80

R = 0.9968

99.68%



Heating

~0.33 Kg ice/L required to cool microfluidizer



(0.33 Kg ice /L )(334 KJ / KJ ice) ( 35 L/min) = heat generated

= 3857.7 KJ/min

=923 Kcal/min or 6.86 Kcal/L





Cooling

923 Kcal/min means 923 Kg change 1ºC every minute

Assuming heat transfer coefficient to be limited by water (and not by resistance of metal)

= 250 BTU / ft2 hºF = 1420 W / m2K

Assume 1 m2 surface area.

6.86 Kcal/L corresponds to 6.86 º C increase since 1L ~ 1 Kg

∆Tln = (15- 6.85) / (15 / 6.86)

= 10.4 ºK



q = UA∆Tln

= (1420 W / m2K)(1 m2)(10.4 ºK)

= 886 KJ/mim

(3857.7 KJ/min) / (886 KJ/min) = 4.35 min









Calculations for Protein A Affinity Column



Determining mass flow rates and volumetric flow rates for Affinity Columns.

conc of mAb coming out of 2nd centrifuge:

0.006mg/mL

want flow rate of 22

mass in per cycle 15829.04

number of cycles 36

mass out of end ion exchange column

16.66667kg/cycle

mass in of end ion exchange column

0.9595% separation

17.54386kg/cycle

mass out of affinity column

17.54386kg/cycle

0.9595% separation

mass in to affinity column

18.46722kg/cycle







81

mass in 18467.22mg/cycle

vol in 3077870mL/cycle

vol in 3077.87L/cycle

time 140min

flow rate 21.98479L/min



assume time is 140min



Now find area using flow rate assume 3:1 diameter to height ratio

But the height should not be much larger than 2 feet.









So we will have 2 protein A affinity columns in series

D= 1.792257m 5.880037feet

H= 0.61m 2.001288feet

V= 1.538935m^3

A(cs) = 2.522844m^2



Determining amount of resin needed, container size, and flow rate

resin needed resin container

12 mg protein/mL resin V= 2000L

mass in 9.71959kg/cycle V= 2m^3

mass in 9719590mg/cycle H = 3D

resin capacity 12mg protein/mL D= 0.921318m

resin vol 809965.8mL H= 2.763953m

resin vol 809.9658L per column

resin load time 60min

resin flow rate 13.49943L/min



Determining amount of elutant needed and flow rate

mass out 17.54386kg/cycle

mass out 17543860mg/cycle

elute volume 2500L

elute volume 2500000mL

conc out: 7.017544mg/mL



Determining pressure drop down column. From Bio-Rad website, the maximum pressure drop

the beads can withstand is 1,000 psi. http://www.bio-

rad.com/B2B/BioRad/product/br_category.jsp?BV_SessionID=@@@@1869198434.114763799

0@@@@&BV_EngineID=ccccaddhkifekmkcfngcfkmdhkkdflm.0&categoryPath=%2fCatalogs

%2fLife+Science+Research%2fChromatography+%7c+Protein+Purification%2fChromatograph

y+Media%2fAffinity&catLevel=5&divName=Process+Separations&loggedIn=true&lang=Engli

sh&country=US&catOID=-31129&isPA=false&serviceLevel=Lit+Request





p v



L k

82

maximum L allowable

max. del p 6896552N/m^2



3

2

dp

k L= 1.983478m

150 (1   ) 2 The height of the column cannot exceed 2 meters





dp= 45um

dp= 0.000045m

assume void fraction is .1

e= 0.1

k= 1.66667E-14m^2

100feet

L= 0.61m 30.48037m

assume u is that of water

u= 0.000798kg/m*s

v=Q/A

Q= 0.010992393m^3/min

Q= 0.000183207m^3/s

A= 2.522844356m^2

v= 7.2619E-05m/s



delta p = 2120970kg/(m*(s^2))

2120970N/(m^2)

307.54065psi



Cost Calculations





Tris-HCl buffer needed

180000L per year

from bio-rad

1 L = $22 for 1.5M

so dilute 1.5M to .1M by 15 dilutions to get 1L of .1M = $1.47

cost of 180000 L $ 264,000.00

1 L = $1.47

Sodium Citrate elutant needed

180000L per year



5 kg = 214british pounds

Assume density of 1 kg/m^3

5 m^3 = 403.9464US Dollars

5 m^3= 5000L



price 0.0807893$/L



cost of 180000 L $14,542.07









83

Cost of 2,000 L protein A container

V (in gallons)

Cost =375*V^.51

from page 556 in Seider



V= 2000L

V= 528.40159gallons

cost $9,177.89



cost of affinity column shell

V= 1540L

V= 406.86922gallons



cost $8,032.55 per column



Anion Exchange Column Calculations



Anion Exchange Column

Total Mab 600kg Mab

total cycle 36cycles

mab/cycle 16.66666667mass mout

% recovery 0.95

mass in 17.54385965kg Mab

Find flow rates and bed volumn dimensions

For two column

time

assume mAb conc in 7.018mg/mL

mass in 8.771929825kg/cycle

mass in 8771929.825mg/cycle

Volume in 1249918.755mL/cycle

Volume in 1249.918755L/cycle

time 55min

flow rate 22.72579555L/min



For two column



Volume per column 624.9593776L

Assume 3:1 diameter to height 0.624959378m^3

Diameter 1.336475661m OR 4.276722ft

Height 0.445491887m OR 1.425574ft

Flow rate per column 11.36289778L/min



Pressure drop - Use Darcy Law k 44.03

pressure drop 687.15psi



mass out 8.333333333kg

elution vol 3CV 1874.878133L

1874878.133mL

conc. Out 4.444733333mg/mL









84

Find resin volumne

Use binding capacity 30mgAb/ml

know need 16.67kg of Ab

so calculated ml of resin needed 555666.67mL









85