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.. jf-T'ar t Jeet Hoac A'v" M ’ Jt ‘ >4r 5' J ’ t- ' --i ~0l AU( 06ur

A NOVI [ I IUROLS B ID BIORI A( t o r io r m a m m a l i a n

(III (II II RI

DINSI k CATION

Presented in Partial lullillment ol the Requirements lor

the I )eriee I )oclnr n| Philosnpln in (In' ( itadualc

School ol ilu‘ Ohio Stale Cmvrrsiiy

In

llui /.Ini, ILS , M S . * 4 • ± *

I Ik- Oilm Stair I 111\ m it\

]00s

I )isser tation ( 'ommillee Approved lyv

I)i Sluing Tian Yanjj l)r Jell ivy.I Chalmers

I )r Joy I* Late Advisvi

Dr Douglas A Kmss I )eparttiienl ol Chemical P.nemcenne UMI Number: 9S3410.1

UMI Microform ') V) 1 I 0 f Copyriqht I'lOh, by UMf (Company. M I riqhts rrsprvml.

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UMI 10 0 North Zeeb Hoad Ann Artior , MI 19 10! In

All Who ha\e Supported Me alonp the Wav

n At kNOYYUHHiMKNT

I expiess ni\ Miiu'ir appiecialion and prat 11ude to m\ advise: I >i Nltanp han

Yarn: loi his oweItfnl puid.mce and i>uIs:.irh1 1nu msiphl thioughoul m\ icseaic It Mis dedication and wisdom h;is led lo much improvement m tins research

Thanks po to other members ol mv advisory oiinniillo'. I )i s l)ouplas kmss, Jo\

Hate. and Jcllrev ( dialmcrs lot iluai piecious puidame, assistance and comments in teinis ol then special knowledge til mammalian eell cultuie. analytical technkpies, and experimental design Ilk- common ellorts made In the12 ‘ ‘a committee has boon a ma|oi icaxon loi the (.puck accomplishment ot tins i esc and)

I learned tin1 general methodology ot mammalian coll c 20 1 1 1 out l)i s Sam

HI ack. ( iihK Baldwin, and Jov Hate I lie knowledge that I attained in then course helped me all through tins woik

Ms Leslie S Jones and Ms Margaiet (Inane aie piatetulk acknowledged tm iheii kindness ol piovidirip the luteal tells, anak/mp the pi opestei one. and suppkmp mlormalion and knowledge ol the cells Ms Doina Lapusan is smceivly thanked loi piepaimg the li'mohlasl cells and helping me analyze prostaglandin Ik and cell density

Mr Jm Miao /lianp is also acknowledged loi his knowledge ol computet modeling

Thanks also go to people ami lahoialones supplying cells aiu! cell lines. Dr

('indy Baldwin loi hyhiidoma cell ILA Th Dr Richard Mollen.son loi hyhndoma cell

111)21: I )i. I )outtlas Kniss loi ! ibrohlasl N Hi V[ V and l)i Joy Hate lot bovine luteal

in svlls l)i Jottlev ( linlms'is is also as knoyvls'ihjs'il lor ktndlv providm p ss 'II s'ultuis' lavihlk's slminp all ol tils' s'\ps'i iitk'iits

I was dsvplv movvsl hv aiul llianklul to mv colk'apus's. I )r Jvh WVn Ys-n. i■ I!s'n

Silva, Yanp Mum I.o. Yan lluaim. 1 lops' Slum. Xu VYs'i Jin. aiul Yi I is' loi maknm a u i uuls'i I ill working s'in n onnis'nt ansi h s' i n vt paik'iit with n w siuimp mv livs' wars in tins

L’lOlip

I'o all mv lik'inls. thank \m i loi yiviny m r all tils' Mippoits. company. aiul all the pai tis's shir iiil! ills' wos'kt'iuls

T. * m\ lovuijj latlk'i and modis'i, m\ Inst Isas hs’i s aiul advisors to ills' Ids' I love you \s'i \ ninch ansi thank you lot L’ivnm ms’ so mush that will always Ik- pis's'ious in my winds' Ills'

l ast, to my yule aiul sls's'ps'st loss', .loan YVu. tils' ps'ison yvho has shanyk'sl my yvhols- Ids', dia.nk you loi yom minn'iiss' low. utisis’i slaiiiimik help ansi s'lis'ouiays'mi'iit. as yy s 11 as sw s'ryday company that maks's my Ids' so ins' an i n t: I ul aikl yvonsis'i I ul

IV MIA

Nil’, m ilv i .10. I Of,: Hi.[II S 1K1111J IK11. ( IlllKl ll>M) ION-1 H S Bioclk'mical I'njjiiuvi ihl' liaO ( lima 1 'nivi’i sii\ ot 1 acliunions Sliaiijjli.il. C 'Inn.i

]0‘U |OSO ('lima 1 n v 11a >nn ic nt a I ISoil’l 1 11*n Ajji'iky kt'si'aidi Inst it UK' in Shanghai. Shanghai. ( hma

I 000 [00: M S ( 'lu-mical I'liLMikvr inti. Ilk' ( Muo Stall' I 'imvrsilv . ( 'nlumhuv ( Muo. I ■ S A

ou 'scmi (iiailualr Ri'si'arvh A.wuiati' ( Ik'mical laijjiikvnnji. I hi' < )hm Stati' I anvor sit v, C nlumluis. ( )hii i. I ! S. A

\ P I 1 111 ,!C A T IO V S

Nun.l I). If. / 1) 11 (i / I a v . aiul R ( Sin ( 11 > S N i I icatmcnt ol R Icavlnn and i>\ciML’ Wash- W'ak- 1 h\ Riohmival Ovulation ami ('henikal I’hvsu.d I’u'cipilalion icchnoloav Nliaindiai Iin iionmental Sciences. Vol 7. No !. 77 :s

/In i, II. (IONS) l iih/alion ol Immohili/mi ln/smcs m Treatment oi the Wasiv Walet liom Slaich Riocessinjj Industiv Nliamjhat Invironim'iilal Sciences. Vol 7, No S. is s

Yanm S [ . I ( lariL’ aiul II. Z lm ( 1 t>i>7! > A N o w l l ei mentation Rroccvs loi (’alcium Miuiii'Miim Acetate iCMAl Production liom ('hec.se Wlicv A[>pI Riochcm Rnxechnol . i )/^s Sno S(sJ

I Yaim. S. I II. Z h u , V I* I a'wi.s. and I. ('. I a n I I * >2 ). Calcium Vlajjncsium Acetate (( \1 A ) Production liom Wlu'\ Permeate: Process and Pconomic AnaK.sh. Resources. Consei cation and Rcc\clms: Vol 7: IM 700

S I A aim . II. Z h u aiul Y la ( l*>‘)4i Continuous Ptopionatc Pioduction horn W he\ Permeate I sum a Novel 1 nhious Red Inoivacloi Rioicclmol N Rioeim \ ol O II 71 I I hi

FII I DOKSTl I)V

\1 u 11 >t lucid. Chemical I'immecnnjj

Mmoi I icld Riochcinical limmccrme Rinchcmisti \ I Alii i: O K O M KNTS

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I IS I ( )l ! K II IMS M il

1. IV I R< >1)1 ( I ION 1

1.1 M am m alian ( Yll ( allure I 1.2 Mammalian ( VII Rinreartms 1. i < >h jectives I

II. U T K K A T l RI, KI-2V1KW 7

2.1 Introduction . 7 2.2 Suspension ( ell Biorcaeturs X 2.4 Iinmohiliied ( ell liioreactors ]() 2A I Micmt'arnor Bioivac'ois . i() 2 M2 Hollow I ilvr liioicacUns ...... II 2 i. 1 I i\i.\l lied liimeaelois I ( 2.4 Mode ol’Operation: Batch, l-ed Batch and Perfusion I X 2.5 f ibrohlast ( ell ( 'allure I (. 2.6 Primary I.uteal < ell ( ulture Is

V I! 2.7 IlYhridnina Cell Culture 2 7 ] Monoclonal Anlibodx i’rodiiclioti 2 7 11 Suspension ( Clime 2 7 12 Immobilized ( ell ( Clime . 2.S Hvhridoina ( ell kinetics and Simulation of the Fibrous lied liioreactor

III.] IliRORLAST CKI.C (T I Tl RK

VI Introduction

3.2 Materials and Methods i 2 I < VII I .me i 2 2 Medium '2 'I Vic run i ul i on ol < )plmi.il ( once ill ial ion ol I’M A i 2 I ( Clinic ai t llasks i 2 ‘s ( 'ultlito in Miciocamei Hioicactois i 2 C (' ul tins' m a I ihious Hod liioieaeloi V2.C I Seleelioii ol I’ackiiiL’ Matenals in die BioieaUoi i 2 (v2 Well Mixed Biorenclni t 2 I' t I’hiu Flou liioieaeloi { 2 7 liioieaeloi Stall tip V2 b PC i I - ; l’i odiiclion m the liioieaeloi undei Hatch Condition V2 o PC d ’ l*i oduelion m Him caclor under ( 'on11ihious ( ’onditu i 2 10 Analytical Melliod.s V2 10 I {'clI l)ensii\ in Sialic Fla-ks I 2 I 0 2 (ilueo.se and I .actale ( 'oneenlralion...... 1 2 10 7 (rlutanune ( 'oncentr alion ...... V2.10.1 l'i 'oslaulandm I..1 I PC il:; I Coiicen trail on 02 10 3 ( VII I Vnsiis m Mieioca11 lei liioieaeloi s t 2 lOOCell lViisi(\ hi Fibious lied Bimcaetois <2 10 7 Seaniiini' Meelion Micioscopx (SFMi

3.3 Results and Discussions silt )plimal < oiu cnlialion ol I’M A ' s 2 K 111elis s ol Slain. I Mask ( Clime s VI kinetics ol Mieioeamei Spmnei Mask ( Cllme 7 3.3 Selection ol liioieaeloi Backing Maleiial.s ...... 7 0S Kinetics m a Batch Mode Fibrous Bed Bioreaeioi ...... 3 3C Kineiie in well Mixed liioieaeloi...... i i 7 Kmeiics in Plup Flow liioieaeloi 3.7 X Comparison ol Various ( 'illlure S\stems 3 3.0 Kinetics ol Suh.strale.s. Metabolite and 1'ioduel ( ( 10 Scanning Flection Micrographs VC 1 1 Conclusions......

IV PRIMARY IT ITAL ( FIT Cl I.Tl'KF

4.1 Introduction 4.2 Materials and Methods 7 4 ■4 2 1 I ^is.socialitwi ot ( ’orpus I aiteum <( 'I. I 7.4 •4 2.2. ( ullure Medium 7-4 ■4 2.4 C 'ullivalum m Static ( 'ulture I I;isk^ 7 1 I 2 1 ( diltivatioit in a Hbrous Bed Bioieactoi 7s ■4.2 4.1 Bioreactor Construction...... 74 ■4.2.4 2 Select ion ol Rackinu Mater ials ...... 70 4 2 4 .4 (, 'ell Imiimhili/alion . 7 7 4 2 4 4 ( dilti valion under Balt h Mde 77 4 2 4 s (’u 1 ri\ alii'ii undei Ted Batch Mode 7N 4 2 4 o ( 'ultuie undei ( 'oiitnuious Mode 7S

I 2 s Ana! \ I icail Mel In nit 7N 4 2 4 1 ( til I)(.'iini(\ m the I Husks 7N 4 2 4 J (ilucose jnd I actate 7X 4 2 4 \ Ri ocvslei one .... /0 4 2 s | Scaniiim.1 Idcclion Micioscop\ (SI M) /*>

4.4 Results NO 4 4 1 I'ropcsleione Production h\ I 11 Siimulalmu NO 14 2 ( 'ultivalion in Slain: Husk NO 4 4 4 ( 'ultivalion in luhtous Bed Bioieactoi N1 4 4 4. | Selection ol lubmu.s Materials ...... XI 4 4 4 2 Cultivation ol 1 .utcal Cells in lihrnus Bed Bioreacioi N2 4 4 4 4 Cell VA4C , ' . ! ipies and AllachmentConditions . . N2

4.4 Disc ussions N 4 4 4 1 ( 'ell Interaction . N 4 4 4 2 ( 'ell Moi pliolop v X4 4 I 4 Nuinent I'aetms X4 4 4 4 Mass Ttanslci Rale NO

IIVBRIDOMA < I I 1. CCCTl RI 102

. I Intr oduction I 02 .2 Materials and Methods 1 0‘s s .2 1 Cell lane ...... I OS 4.2.2. Medium ...... 100 V2 .4 Batch Cullure studies...... 100 4.2 4 1 Id lect ol Cell Seeding I )ensit\ ...... 100 4 2 4 2 Static [Mask 100 4 2 4 4 Spinnei Masks 100 4 2 4 I ihious Bed Bioreactor ...... 107 4.2.4.1 Selection ol Racking Materials loi Bioieactoi 107 .4 2.4.2 Bioieactoi Construction ...... 107 4 2.4 Bioieactoi Start u p ...... ION 4 2.0 Kinetics Studies ...... 100 4.2 0 I Well Mixed 100 4 2 0 2 Blue I low 110

i \ X 2 7 I .one In in Stability IB) X 2 S Bed Hatch Kinetics 110 X 2 o Analytical Methods ...... X 20 1 Cell Densils ami Viability x 2 0 2 C11 uc use and I.aclale Concentrations ...... X 2 0 A Monoclonal Antibody ( ’onccnlt ation . I ll X 2 0 4 ( ihitamine ( 'onccntralion 112 X.2 OS (V]l Density in Bibious Bed Bioieactoi X 2 *> (' Scanning election Micioscopy (SB.M) . X 2 0 7 ( 'onh>c. il M k aoscojn

5.4 Results and Discussions X ^ I I -1 led ol ( ’ell Sc 'dm;! Density 11^ x B2 Kinetics to Hatch ( 30 ’’ s ol I h In id on i a ( ells 111 X X J | Static 1 Masks 111 x < 2 2 Spinnei Biask 1 I s’ x < ^ Select ior, ot 1 hi 1 ons Matenals toi Cell hnmohi I i/at ion I I x X < 4 Kinetics ot 1 Ivt'i idomas m lihnuis Bed Bioieactoi I IX X 7,1.1 Well M i\;d ( ultuie ...... I I ' X < -1.2 Blue Blow 5 I to C | t Bed Batch ( ultuie I I 0 X A X | 11 ii tJ lei m Stability ot the Bibious Bed Bioieactoi s d h Cell Density and Viability in Bihmus Bed Bioieactoi s IIS x 1 7 ( 'onipanson ol MAh production in various ( ’ultuieSystems . [ I ,x X S SBY1 ol flyhridoma Cells in a Bihious Bed Bioieactoi I ]0

5.4 ( (inclusions 120

VI HYKKIDOMA ( I I I. KINKTICS AM) SIMM,ATION Ol Till I IliROl S BIC19 At I OR IK

6.1 Introduction llo

6.2 Kinetic s ol 1'2825 1 7 ‘ a Cell Culture IIS 0 2 ! K i nr (k Mode I IIS 0 2 2 R.u amrtri B.sl m ml ion .. . 1X1) 0 2 2 1 ( irowth and Pioducl Viekt 1X0 0.2.2.2 Spectlie MAb Bioduciivitv...... I XI (.2 2 A Rate ( "onstanis ...... I X |

6..A Mathematical Model Cor Kihrous lied liioreactor 1X2 0 A I Pressure l)io[i and Shear Boicr in the Fibrous Bed |S2 6. \ 2 Mass I rranster Coelticient in die Bibious Bed ...... t X A 0 d.d Mass Tianslcr with reaction tn the Fibrous Bed I X4 (rd-I I.k|unl Blow Velocity...... IXX

6.4 Simulation of Fihrouss lied Bioreactor 1X0 (i 4 I Concentration ( iradienls in B’lbrous Bed 1 X6 0 1 2 ( 'ompanson ot Results liom BAprrimrnls and Model Simulation I X7

y (' 4 4 Id lee 1 ol I'oiosity ul I Ik1 lihmus Matrix on MAhhr odiicl ion I TS fvT I 1 .1 loci ol ihe Thickness til (he hthcr Mams ...... ISX h ■! S II I eel ol the Thickness Ratio ol Matrix to ( Tip, a/h I S1) h i h f iled ol Recu cuialmn Rale on MAh i’mduclion . IhO lUliTMM.K\RHY I7S a i *i *i-:m »k i s

A .A compulei pioijiam lot simulation ol ihe lihious hod hioieator I Kh

li Id ISA hu MAh deletmmattou l(,(>

( ' ( ilulainine Analysis 1 h I

I ) ( ell I Vnsilx I )eler m malioii ! 7

\ I LIST OK TAIiLI S

I Mil I P \(

2 I f ' 1.is^11main>n (>! mammalian veil hmicav ini aCCnldllie In in.mini [\|V 2{>

2 2 Nui'lace \ nlLimn i al in 1c m 1 I nl \ at mus melhnds in cultui v aiichnra;je dependent noils ill

2 * ( 'nmpai isnn n| suspensmn and immnhih/ed cell culluu 1 echiikiues

' I ( nmpai isnn nl dll leienl systems ! nil ihi nhlasl cell culliiie M

d I ('nmpai isnn nl pi neeslerune respunse In luteal cells In 1,11 05 ‘ in various culturim.' systems...... N7

s ! \1 Ah producliviIs and plucose ami elulamine vniisuinplinn lain in lliv Ind hatch cultuie

^ 2 ( i m i|iai isn 11 nl \ annus s\ stems Ini In h iid n m a cc HI) 2d culture

s ( (ddl viahililv and distribution m die libious bed bioieactoi

h I Parameters and then values Ini die kinetics mndel Ini liyhridnma cell cullm e ...... Ihl

2 1 lie parameters and then values used m compulei simulation olthe lihrous bed hmreactni in]

I The proceduie nl leapents and sample addition IP's

\ 11 LIST OL I !Ol RLN l(il Hi HALT

I i lie C\cln n \ \ L'C llase alhl ItpnW LVIia.se putlmaV nl aias liklomc acid metabolism T1

2 2 I he metabolism nl sliimilalmn nl lulemt/mp hormone and lipoproteins In luteal cells aik! the pioduction aiui secretion nl propesterone O

2 2 Schematic diapiam nl the pn»tinet 11>n nl hvbruionia cells secH'lmp monoclonal antibodies...... VI

d -1 Schematic diapiam o! hsbiidoma cell metabolic patlma\ ^

s I The spnalK wound stMicluic nl libimis packim* m 1 hi- im iieaelni ss

s d Selienialie diapium nl m il mixed libmii' bed bimeai. ha b a I ibi nbhna i el I eulliue v,

2.2 Seheinalie diagram nl plug llovs libious bed bmreaetni Ini Itbrohlast cell culture ......

Vd I ^'termination nl opmnal I’MA concentration Ini cell slmiulalinn ...... vs

*S Kni'-tics nl I ibn blast cells Nil I U ^ m static Husks sx

xtti Wi Kinetics ol I il'ioblast cells NIII V| N...on micmcaiiici.x ......

V7 SLM pictures ol lihrohlusl colls cultuicd wilh v;inmis fibrous niatcruils P<)

V.K Kinetics ot 1hioblast cells N lll VM m the tihious bed bioieactoi at batch mode (> 1

' d Kinetics ot libroblnst cells Nlll V[ i m the tibions Inal bioieactoi at well mi xed mode << 1

s 10 Kinetics ol 1 hi oblasl cells Nlll <1 i in the tihrons bed bioieactoi at pin l; I low nnuie ('7

i I 1 ( 'ompai icon ol IJ( rI ; projuelion bv libioblasl cells ill carious culiine s\s|ems (id

■1.12 Lactate yield from plucose in various systems ...... (is

s 1 s Li ostaidandm L; yield liom cliicose in static I lasks and microcamei spmnei lla.sk ...... M

i 1 i Li ostaplatklin I n ucld liom iducose m I ihious bed himeactoi . (id

s I'' Liostaplaiidm I ; \ieid liom jdulamme in hbious bed bmieai loi (> )

VIP Scanning electron microscopy photos o! hhrobla.sl cells N lll V| V in I ihious matrix ...... PS d I Schematic diagram ot well mixed libious bed bioreaclor lor luteal cell culture...... NN

■I 2 Lropesteione production by luteal cells m lesponse to LI I addition in the t ihious Ivd bioieactoi ...... Isd

x i \ -I ^ The piotjesterons'secielinn ol lutealcell cullutcd in slain: lla.sk ...... 1

1-1 rifccI ot medium change Ireijuency on proL’oslcrone p:*hIik 11<<11 ...... (,iI

■1 ^ SI M pliolos ol luteal colls atlaehment to various tihious nialonals .. . . i

1 S 1’ioL’ostoiono pi ndiiclion h\ lutoal cells culluicd m I ll'l OUS he'll tliOUMOlOl ‘*2

1 7 ( ’ompansun ol pi otjesici one pioduclioii h\ luloai cells io I II ill vai ions s \stems......

1 is SI At [lhoios ot luloa! ve il moi pholoeies aiul altaolmionl i oiul111oii in stalls' llasks lM

1 SI M photos oj luloai ooll moiphohmies and atlaslmioiil oondilion in tihious bed liioieaeloi ‘J's

V I Tils' spirally wound sliucluic ol lihrous pashms toi llio tihious hod hioioaotoi Id<

^ d A schematic diasjiam ol well mixed libious hod bioieactoi s\ sU'iii ! d 1

' A schematic diapiaiii ol plus' llou tihious hod hioicactot ss sIo111...... 1 ds

Ad Idlest ol dillcrcnt cell secdmij concentration on cell jjrowth . Idh

7 S Kinetics ol MAh production hy hvhridonia colls culluicd ill stalls I tlask ...... 1 d7

Ab I.moai plot ol (a) coll density vs. fjluco.so concciiliatmn and (hi lactate concenlr at ion vs fducose conconlratioii tor T Ila.sk cullurc . I dS

\ v 7 7 Spt'L 11 ic M Ah productivity versus specilic prowth rate I 71)

AS Linear plot MAh concentration veisus cell density in llie sialic I Mask 1 .H>

7 c) I ancai plot ul (a) MAh \ s plucose concen1 1 at ion and ( h) MAh concent ini ion \ s elulamine concent lation I hi

7 ID K i lie I ics 111 h\ hndoina cell III > 7 1 cuhui ed m spmnei 1 task I ' I

7 I I I.meal plol ul lactale com.enilalion veisus eluc o

7.17 S IM ol In hi idoma cells immohiti/ed in \ at ions I ihious main \ materials I * *

7 I ^ Kinetics ol livhtulonia cell 111) 7 1 m hhrous bed hioteacloi under well mixed conditions 17-1

7 11 K met ics ol In hndoma cell III) 7 1 in ! i hious bed hioi eacloi undei [ilup How conditions I A

V I s Kinetics ol h vhi idoma cel I 111) 7 1 i n I ibi ous bed bioieactoi undei plup llou conditions I In

■s !h K metics ol i a i lactale pi oduetmn I nun plucose. t h) MAh production liom plueo.se. and (ci MAh pioduclion liom elutaiiime m libious led biorcm loi I W

‘s. 17 Comparison id MAh production in various cultuie systems 1 h)

7 tl) I.one term stability test ol the libious bed bioieactoi I V) s.70 NLM photos ol hyhiidoma cells in libious bed bioieactoi I K)

“s 71 Conlocal microscopv photos ot stained hvbrtdoma cells at dilterent depth ol libious matns 111

\ vi b I Se‘IK'in,i11c clia^iam <»l mass tianstsi rsyimam Iihnuis bad system lb.1

(i 2 I lia biu-ai ploi ol si'll ilen.siiy vm su-, ylucoss-coiKanlration . HA

b.d Tha lineal plot ol las'tats' sonsantralion versus ylus'o.se sonsanlralion ...... Ibd

n I 11k- ii’I.iiuinsIii11 heivuvn s|H'l 11is MAb poulium 11\ aiul spivilk1 yiowili tali' Ib l

ft s 11 \ hriilomu ivli slan-uls s ban ye slut me stalls Husk suit ut a Ibl

b b ( ilusose sonsenlialioii slianya dm any skills' lla.sk sultuie ...... I bs

(i 7 1 astati'soiuviiliulion slianya shinny slain tlask sultuie lbs

bN Ksdalioiiship ol piessuie slrop and meslium velocity Ibb

bb lisa * I) MAb concentration piolih' m llu- hhei mains ...... Ibb

till) Ilia M ) ( ibis osa s oiuanl lalmn [u ol i Is- m ilia I dial mad a s lb/

n i l t ha s I) I as’lala aonaa im al ion pi olds- in i ha 1 1 luo mains ! b7

b id ( ompai i si m ol a spa 11 mania I data and modal Simula I ion ol hvhrkloma culture in pi lip I low tihious bail bioraaaloi I b,X

b id Idlaalol porosity on VlAb piosluationand ylucose consumption Ibb

(vl d Id lust ol basl poiositv on volumetric Now rata ratio ansi spaadic MAh production ...... Ibb b ! s I tiesiot thickness ol 1 ibei mains on M Ab pioslualion anil ylueosa eonsentiation...... 170

s\ II (i If' Id led ol the recnculation late on MAh ptoduclion 170 h 17 Idled o! thickness ratio ol mains lo pup on MAh production and plucose consumption 171 fi IS Idled ol thickness ratio ot matrix to pap on specific MAh production and voluntelric I loss late ...... 171

B i Slandaid cm \ e lot MAh delei mutation |0s

( 7 1 lie slandaid cuke loi plutannne mcasui enient loo

I) I Slandaid dir\e ol cell density determination hv protein assay ...... 100

w in (II AFTER I

INTRODUCTION

1.1 Mammalian ( ell Culture

It has hivn a long time since people started to culture animal cells in vitro As early as m l‘K>7. Rv ess Hamison successtullv grew embryonic nerve cells hv the hanging drop technique loi several weeks In the late HHO.s, alter haulers discovered that viruses could he propagated in cell culture and used as vaccines, large scale animal cell culture lirsi came into the picture Since the ld'SO.s. animal cell culture has undergone many considerable developments and more and more cell lines have been attained Irom dillerrnt animals and dillerent organs I'he dramatic advance ol recombinant DNA technology in the ld70s made it possible to culture mammalian cells to produce protein products. Now, a range ot recombinant mammalian cell cultures hav ■ been developed w ith the capability ot producing various selected products

Applications ot the products produced b\ culturing mammalian cells aie mcieasmg in many aieas M o.st ot (he protein products aie used as ptiai maceutical. diagnostic, therapeutical, or veterinary products and have good and constantly increasing markets Many culture systems have been developed lor cultivation ot mammalian cells and harvesting the products However, there are still many barriers to be overcome loi conventional mammalian cell culture, such as low cell density and product concentration, shear stress damage, mass transier limitation, and poor long-term stability, bor every system, a carelul study is needed on the characteristics ot the cell line, product secretion. medium formulation. the culturing conditions, the properties ol aeration, shear loree and

mass transfer, and scale up potentials.

I'here are two types ol mammalian cells: anchorage-dependent cells (cells that icquue a surface to at 1 ach to lor growth), and suspension or anchorage-independent cells

(cells that do not require a surlace to attach to lor growth) I he traelitional method lor the laigc scale produclioti ol anchorage depeiulent cells uses the inside suilace area ol rotating bottles lor cel) growth Since the surlace area lor cell attachment is only a small percentage ol the total bottle volume and a large number ol bottles are needed to produce even a small number ol cells, this method becomes cumbersome and expensive for producing a large quantity ol cells Thus, the tatio of the surface area available lor cell grouth to the total culture volume must be increased lor success!ul large scale mammalian cell cultivation On the other hand, anchorage-independent cells can generally be cultivated like microorganisms by conventional lermenters However, because mammalian cells do not have a protective cell wall, they are much more sensitive to shear forces than microbial cells Thus, shear reduction has been the goal that catalyzed the development ol new techniques lor cultivating suspension cells

1.2 Mammalian ( ell Bioreactors

The bioreactor design lor mammalian cell cultures is an area undergoing rapid development. The first industrial animal cell bioreactor process was designed in the mid

IbSOs tit produce polio vaccine from African green monkey kidney cells grown as an adherent culture in either static flasks or rolling bottles. I'he scale-up ol the process simply involved increasing in the number of flasks and bottles

To increase the ratio of the surface area to the total culture volume for anchorage dependent cell culture. Van We/el developed the microcarrier technique in Id67 The microcarriers permitted the culture of anchorage dependent cells in relatively standard stirred lanks, making cell culture production similar in convenlional mk'rmii'jianiMii

lermenlation Because some problems existed111 tin.' application ol this technique lo routine production, many facilities maintained the use ol flasks and roller hollies into the

IdKOs The ma|or technological innovation m ll)7()s was the development ol hollow liber hioteaciots B\ using tins technique, dillusion ol nutrients occurs across the memhiane ol (he hirers, hut the cells immobilized in and around the hollow libers aie

[uevented Irom being washed out Although this approach received much early attention, it was not wideh used on a manufacturing scale until recentlv. Since then, mam othei novel hioreactors have also been developed to increase the cell density anti pioduct productivity

In the attempt to reduce the shear toice to which the anchorage-independent cells are exposed, several diflerenl modifications on bioreactor agitators and air sparging techniques were developed. Vibromixer, marine propeller and other agitator designs were developed to reduce the shear iorces by producing a more gentle agitation. Changes in air sparging .strategies included creating an artificial cell wall analogous to that ol microbial cells, encapsulating the cells inside a man made membrane, adding surlace active agents, ot using bubble liee aeration techniques to improve cell viability undei sparging condition

During the last ten years, there has been a radical change in the types ol mammalian cell bioreactor systems being designed. Large scale bioreactors were successfully constructed and operated Because rnosi industrial scale cultivation systems were batch bioreactors, the cell density and productivity wane still low Some novel hioreactors could achieve high cell density and product productivity, however, most of them suffered from poor mass transfer, reactor louling. and could not maintain a good long ie»rn operation stability Iliere are several challenges remaining in animal cell bioreactor design: (I) cell culture production is still expensive in most cases, (2) most 4 types ot ihe bioreactors lace scale up problems, < T there is a trade ot( between high cells

density versus long term stability. and (4) there is not much information on mass transfer

and kinetics in most novel bioreactors

fn this research, an innovative fibrous bed bioreactor was developed toi

mammalian cell culture This continuous immobilized cell bioreactor overcomes mans

dinwbacks ot the son vent mnal bioreactor by using a unique, spirally wound packing

structure to provides the properties ot high cell density, good mass transfer, a three dimensional environment toi cell growth, and good long term .stability Different cell and ceil lines (luteal, tihrohlast. and hvbndoma cells) were cultured in the fibrous bed bioreactor. and the efficiency, properties and operation parameters ol the bioreactor were evaluated experimentally and by computer modeling

1.3 Objectives

By cultivation ot different mammalian cell m the fibrous bed bioreactor, the following tasks were performed

i I i bibmblaxl cells N lll DT from mouse are used as the representative ot anchoiage

dependent cell hues The cells were cultured m stalls I flasks, microcamei

bioreactors. and the fibrous bed bioreactor. Prostaglandin ID (P(ibD). stimulated

by phorbo! 12-myristale 1 V acetate (PMA), and cell density were monitored

during the experiments to assess the efficiency of the various .systems Scanning

electron microscope (SKM) photos were used to study the cell morphologies and

immobilization conditions in the fibrous bed bioreactor The kinetics of nutrient

use and product formation were also analyzed Hybndoma cell is a typical anchorage independent cell lino In this study, mouse

hvhndoma coll III) 24 producing monoclonal antibodies lg(12h woro culturod in

static T flasks, spinnoi flasks, and tho lihrous hod bioreaclor The productivity of

monoclonal antibodies, tho consumption rate of medium, and the coll growth rale

were monitored during the experiments Mathematical models won’ developed to desenbo tho diftoiont cultuto systems Die hai'vest o| monoclonal antibodies and

the final cell density in different systems were the criteria foi the evaluation of the hioreactor and other culture systems Long term stability of monoclonal antibody

production by hybndoma cells and fed batch mode bioreactor were also studied

Besides the efficiency studies, the fibrous bed hioreactor was also used as a cell culture apparatus to study primary bovine luteal cells The bioreactor supplied a three dimensional culturing environment that was close to in vivo conditions, allowing study of the physiological and morphological properties of luteal cells

This study mainly focused on the progesterone production hv luteal cells m response to the stimulation by lutemi/mg hormonetill) in various culture environments, and the phenomena and effect oi cell migiation and morphology changes on progesterone production during long termin vtlto culturing

Based on the kinetics of hyhridoma cell metaholism and the structure of the fibrous bed hioreactoi, a mathematical model coupled with mass transfer and cell kinetics was developed and simulated the culture processes in the fibrous bed bioreactor. Ihe effect of retention time, recirculation rate, liber thickness, and lihrous bed porosity on MAh production in the fibrous bed were studied The model developed was helpful tor understanding the mechanisms of cell growth IS and MAh production, loi optimi/im.' tlit' design and operation ol the bioTcacloi, and lot supplying useful tnlomialton tor hioreactor .scale up CHAPTER II

LITERATURE REVIEW

2.1 Introduction

Mammalian cells air surrounded In a phospholipid bilavei membrane, which

contain1' emhcdilcd cn/vines aiul smiclural proteins, that mediates comniumcadon

between llie cell ami the environment Mammalian cells lack an outer wall ami, as a

result, are highly sensitive to environmental stimuli, such as shear stress, osmotic s lunges. pH. ansi nutiient changes (Rrokop atul Rosenberg, IdXd). According to

Lambert, et al ( ll>X7>. the following factors shouhi be taken into consideration when culturing animal cells in vitro ( I ) mammalian cells are more likely to be damaged b\ mechanical and shear iorce.s, (2) rich and undefined media are required tor mammalian cell culture, and mammalian cells are also iclatively sensitive to impurities in the water supply and medium components, i t) mammalian cells have a iclativelv long doubling time, (d i mammalian sells are obligate aerobes, so oxygen supplv and transiei rale shouhi be c otisideied, (^ ) m.immaliaii si'll giowth bscomes limiled al relatively low cell concentration by I actors including accumulation ol metabolic products (Butler and Spier,

1^X4). and (6) fundamental quantitative information on kinetics of product formation, nutrient utilization, and generation of inhibitory metabolites is limited.

There are two types of mammalian cells: anchorage-dependent cells (cells that icquire a surface to attach for growth), and suspension or anchorage-independent cells

7 8 (cells that do not require a surface to attach to lor growth) Numerous studies have been

done on using modified or novel hioreactors to achieve high cell density ami product

productivity Classification of mammalian cell bioreactors according to reactor type is

summan/ed in I able 2 1

2.2 Suspension Cell Hiorcadors

Most turuot cell lines, suchas HeLa cells and Ivmphoblastoids, are anchorage

independent l-'ssentiallv. these cells can be grown in equipment sim ilar to that used toi

microbial fermentation. However, because mammalian cells do not have a protective cell

wall, they are much more sensitive to shear forces than microbial cells ((ilackcn, I98h)

Many studies have been done to reduce the shear force generated by agitation,

aeration, and other sources. Different designs ol agitators have been developed to reduce

the shear force from agitation Girard (1977) designed the vihromtxer. which was

composed of a horizontal dok with conical apertures attached to a vertical shaft that

rapidly oscillated up and down, to circulate tho medium in a vertical, rathei than a

horizontal direction, resulting in less sheai lorce required to obtain an adequate dispersion

ol the cells Another agitator, designed In l eder and lo b e it ( 1‘fK h . had loin slowlv

turning flexible sheets that span the depth ol the culture fluid, providing adequate dispersion ol the cells at low' shear rates. Shi and his colleagues (1992) developed a new-

type impeller consisting ol a double screen concentric cylindrical cage impeller that

increased the specific screen area and the convective mass transfer rale through an annular cage. By using a perfused bjoreactor equipped with this impeller to culture hybndoma cells, they reached cell density to 3.4xl()7 cells/ml and the monoclonal antihody concentration reached SI2 mg/1. There are also many other important new designs of hioreactor agitators, such as the marine propeller (Oldshue, 1983), helical and anchoi mixer (Oldshue. 1983), hanging stirrer (DeBruyne and Morgan, 1981). and tilm 0 types (Roubicek and Feres. HJX7). that can provide gentle agitation to reduce cell

damage. These works have demonstrated that better agitator designs were essential it

high cell densities were to be achieved

It is a common observation that direct gas sparging ol the culture medium may he

deleterious to cells grown m tree suspension Sensitivity to gas sparging depends on cell

type and si/e. bubble frequency. super! icial gas velocity, and the form ut cell

propagation lam and Ross X1) reduced cell damage by decreasing the chance ol

direct contact between cells and air hubbies J’hey created an artificial cell wall,

analogous to the cell wall that protected cell membranes trom shear damage by

encapsulating cell suspensions inside a polvmer-reinforced calcium alginate membrane

According to Damon ('o t 1FXI ), the capsule produced can be as small as 5 microns and

can contain as many as 10,000 cells dins technology is commercially available lor large

scale production ol suspension cells, including hybndoma cells

A number ol foam stabilizers and viscosity enhancers were assayed to decrease

the e t fee is ol local hydrodynamic stress Although the underlying reason is not fully

understood, the addition ol surface active agent such as Pluromc. silicon oil, or Pluriol

may impnoe cell viability undei the condition ol sparging (Handaetal 1DK7J Similar

results have been achieved by the addition ot viscosity enhancing agents such as

melhylcellulose or polyvinyl pyrolidone {Kaunger and Scheirer, IMK2), Miltenburger and

David ( 10X0), Tyo and Wang (10X1). and Dodge and Hu ( 10X6) placed wound silicone

rubber tubings in fermenters, resulting in bubble free bioreaclors with decreased contact

between cells and air. The tubing was used for both aeration and pH control, greatly reducing foaming and shear stress caused by direct sparging.

The basic principle of the airlilt fermenter was first described by Le Francois

( 10XS). (Jas mixtures are introduced into the culture from a sparge tube al the base of a central draught tube, which causes a reduction in the bulk density of the liquid in the 1(1 draught tube compared with the outer /one of the vessel setting the culture into

circulation Different means of directing (low outer draught tubes were used and

described by Onken and Weiland (1983). The advantages of the air lift bioreactor are

that it induces less shear stress to the shear-sensitive cells than do stirred tank fermenters

while achieving the necessary mixing and oxygen transfer ((’hoi, et al., 1 90<)) Many

articles have been published describing the performance ol air lift fermenters However,

because of variations m the experimental fermenters' geometry and differences in the

physical properties of the fermentation media, the published data were not consistent,

which makes the realization of performance difficult (Ora/.em. et al.. 1979; Siegel, et al .

1980) To further eliminate cell damage by direct sparging, C'hiou and his colleagues

(1991) developed a cell culture system that combined the features of air-lift and packed

bed and achieved a high density ol’C’HO ce 11s (1.2 x It)7 to 3 lx !07 cells/ml).

2.3 Immobilized Cell Biroreactors

The key to designing a bioreactor lor culturing anchorage dependent cells is to

increase the ratio of the surface area available for cell growth to the total culture volume

Various studies have been done on design of novel bioreactors to meet this need The ratios of surface area to the hioreactor volume of some commonly used bioreaclors are listed in Fable 2 2

2.3.1 M icrocarrier Bioreactors

To increase the ratio of the surface area available for growth to the total culture volume. Van Wezel (1967) developed the concept of microcarriers, which involved the suspension of microscopic particles (DLAL-sephadex A-30) m a slowly stirred liquid medium Such particles offered a large surface per unit culture volume for cell attachment. In 1977, Levine et al. (1977) reported that if the density of the DLAb sephadex microcarriers was optimized, the cell could proliferate to a density of 5xl0(l

cells/ml at rmcrocamcr densities up to 3 g/l.

Since the discovery, a wide range of microcamcrs have been successfully used for

the cell culture of anchorage-dependent cells and many are commercially available.

Among those, dextran, plastic, gelatin, and cellulose are most used materials. Most

microcarriers have been optimi/ed at a diameter of 100 200 pm. so that the available

surface area of 7 x 104 jj m- can support a monolayer of over one hundred cells. Thus, by

suspending 12 6 x It)6 dextran microcarriers in a I liter culture, a surface area of 104 cm '

is provided for cell growth and is equivalent to the surface available in .*>() Roux bottles or

20 roller bottles (Butler, 10X7). Typically, this microcarrier concentration can support the

growth of cells from an initial concentration of 2 x 10s cells/ml to a final yield of 2 xl()6

cells/ml. In addition, the three dimensional microcarriers can provide a framework with

which animal cells can both grow and maintain themselves in a sufficient viable slate to

be able to manufacture product materials in quantities that are economically attractive

The problems facing microcarrier culture are that many studies must be done lor

hioreactor scale up, especially for understanding the relationship among agitation speed,

cell damage, and nucrocarrier settlement, tor lowering the costs of the microcarrier, and

lor improving long term stability ol bioreactors.

2.3.2 Hollow Fiber Bioreactors

The potential of hollow Ober filtration devices for high density cell growth was

first recognized by Knazek et al.(ltf72). Typical configurations of hollow fiber

bioreactors are cylindrical shells and flat sheets The cylindrical shell design consists of

hundreds or thousands of fibers formed into bundles and contained in cartridge housing.

Cells are inoculated on the shell side while medium and oxygen are recycled axially through the lumen of each fiber. Low molecular weight substances such as glucose. vitamins, amino acids, and oxygen arc readily transported across the membrane and

nourish the cells, while the low molecular weight toxic metabolites, such as lactic acid

and ammonia, are transported away from the cells to the outer side of the membrane, and

mlo the recycling medium The cells, which have large molecular weight, are retained in

the capillary space and grow to a high density (Heath, 1 WO)

Moll ow liber bioreaclors have been successfully used for the production of cancer

cells (Hager, et al 1982), viral antigens (McAlcer, 1983), and other products. By culturing hybridoma cells in a hollow' fiber bioreactor, cell density can be increased by

two orders of magnitude and the productivity can be increased by 20 to 40 times of that in static flasks (Klimman and Mckearn. 1981) The modified hollow fiber hioreactor technology being used for the commercial production of cell products is the In Vitron static maintenance reactor (Tolbert et al, 1988) and membrane {Scheirer 1988) which is a sandwich of flat membranes with different molecular weight cut-offs.

In 1972, Kna/ek (1972) had already recogni/ed gradient formation in hollow fiber bioreaclors, especially for dissolved oxygen, as one of the mam problems, which is and he attempted to overcome oxygen limitation by using a mixed-bed hollow fiber bioreactor, but it was not successful In addition to hydrophilic libers transporting nutrients, silicone rubber capillaries were used for oxygen transfer. Some computer models were developed to describe mass transfer, especially oxygen transfer phenomena inside hollow fiber bioreactor, to improve the mass transfer efficiency in the system and therefore, the cell density (Brotherton and Chau, 1990; Piret and Cooney, 1990). By improving the mass transfer rate, Brotherton and Chau (1990) increased the cell density up to 10*-1 ()y cells/ml,

The advantages of hollow fiber dialysis culture system are: (1) in the cartridges, cells can he propagated at densities more than two orders of magnitude higher than in conventional system, yielding concentrated and less impure products, (2) nutrients and 13 oxygen gradients can be kept to a minimum through high mass transfer rate with the

dense capillary network with a large surface area of very thin membranes, and (3) the

hollow fiber bioreactors are suitable for animal and human cells, both anchorage-

dependent and suspension cell types. The disadvantages that keep hollow fiber

bioreactors from being accepted as the most efficient production technique are that: (1)

with various parts and connections the sterility of the hioreactor is hard to maintain, (2)

the high cell densities cause fiber clogging and fouling, and (3) the formation of gradient

within the cartridges limits the scale up of the hioreactor (Tharakan and Chau, 19X6)

2 3 3 Fixed Bed Bioreactors

Many different packing materials have been used with fixed bed bioreactors

suitable for animal cells (Spier, 1985). The most commonly-used substrate is the glass sphere, despite the fact that a sphere has the lowest surface area to volume ratio. A fixed bed matrix has to be a compromise between maximizing the surface area and providing a passage open enough lor medium to be perfused at a rate sufficient to maintain homogeneity throughout the bed without subjecting the cells to damaging shear effects

(Griffiths, 19XX) Looby and Griffiths (19X7) demonstrated that a pore size of over 3pm diameter was needed for maximum growth. Most fixed bed bioreactors are column shaped with support particles on or in which cells arc immobilized. The medium is perfused through the packing materials to supply the nutrients for cells. The packing materials can be divided to two main types: organic and inorganic The organic materials include proteins, polysaccharides, and synthetic polymers, and the latter can be glass, silica, ceramic, fibers, and metals (Ho and Wang, 1991). Burbidge (19X0) reported that a

10 liter fixed bed bioreactor with a packing material as glass beads (2-7 mm diameters) was used to culture human diploid fibroblasts (HDF) and concluded that the behaviors of cells in the glass sphere bioreactor were identical to those in static Basks. A fixed bed. 14 mass culturing system was developed hy Brown (1985). by which a variety of cells were

successfully cultured. A total of 1.3 x 1()'() viable cells have been recovered from the 12

liter bioreactor, and the system has been kept continuously running for over 1 year.

Immobilized cell fixed beds, which allowed differentiation, were also used for obtaining

tissue or even organ-like growth of cells, (Leighton, 1951) lor studying cells in

physiologically stable environments by medium perfusion and as a mass cell culture

technique. Park and Stephanopoulos ( 1993) investigated a porous ceramic particle

packed bed to culture rat pituitary cells to produce insulin, and the cell density reached

2.6x1 ()-s (cell/cm2 of the surface). Chiou and his colleagues (1991) developed a cell

culture system that combined the features of air-lift, fiber-bed, and packed bed with 24-

gm glass fiber and achieved a high density of C'HO cells (1.2 x 107 to 5.1 x 1()7 cells/ml).

New Brunswick Scientific Co., Inc. (Wang et al., 1992; Biomass News, 1993) developed

a packed-bcd hioreactoi (CelliGen) lor which productivity could reach as high as 12-fold

that in static and stirred suspension culture systems.

Fixed bed immobilized cell bioreactors can be used for both large volume cell

culture and the study of cells in three-dimensional environment, because they have many

features desirable in cell culture, i.e. high cell retention and perfusion capabilities, simple

scale up and bubble-free operation, in addition to the potential medium benefits enabled

by the high cell densities (Whiteside, 1981). However, although many packed bed

bioreactors can achieve high cell densities and productivity, they are not used widely in

industries. This is because there is a lack of efficient packing materials, there are intrinsic

problems with gradients and low unit productivity, and fouling and clogging are problems

for long-term operation (Ho and Wang, 1991).

Recently, a new, continuous, fibrous-bed, immobilized cell bioreactor was developed for continuous cultivation of mammalian cells and microorganisms (Yang et al. 1992a, 1992b, 1994). The fibrous bed hioreactor contains a spiral wound fibrous sheet 15 lor cell immobilization, with significant spaces between wound layers to allow good mass

transfer of nutrients and products. The fibrous bed bioreactor provides not only a large

surface area for cell attachment, which results in a high cell density and product

productivity, but also a conducive environment for cell growth in three dimensions,

allowing cells to maintain their native morphology and differentiated properties. The

continuous process using fibrous bed bioreactor has been shown that cell immobilization

in this bioreactor was not irreversible as constant growth of new cells and sloughing off

dead cells occurred within the reactor. The hioreactor is thus self-renewing and

eventually establishes a dynamic steady-state cell population. The structured fibrous bed

allows for good multi-phase flows and provides renewable surface for cell

immobilization. This bioreactor offers long-term stability for continuous operation

without suffering aging, fouling, or degeneration problems, which are normally observed

with conventional packed bed and membrane hioreactor In this work, this novel fibrous-

bed bioreactor was used for dif ferent mammalian cell cultures.

2.4 Mode of Operation : Hatch, Fed-Batch, and Perfusion

Mammalian cells bioreaclors can be operated in three modes: batch, fed-batch, and perfusion. In a batch system, no nutrients are replenished except oxygen. The only parameters that can be controlled are temperature, aeration, and pH. Thus, the cells in batch cultures are subject to a constantly changing environment where nutrients are being depleted while waste metabolites accumulate. Cell growth and/or product formation can be prematurely inhibited due to nutrient limitation and the toxic waste products build-up

(Tovey, 1980)

A better mode of operation may be a fed-batch system that feeds vital components only as needed to the culture, thus obtaining a constant nutrient concentration {Glacken el al., 1983). The fed-batch process requires that methods exist to monitor and control the 16 nutrients in question. More environmental parameters are controllable in this system and

thus, the cell growth and productivity can be better optimized. Although cellular waste

products are still allowed to accumulate in the system, it may be possible to lim it the

accumulation of these waste metabolites by adjusting the frequency of nutrient addition.

Perfusion is the third method for growing mammalian cells in culture. According

to Feder and Tolbert (1985) and Van We/el et al. (1985), in a perfusion system, both

nutrient and waste product concentration can be controlled by varying the dilution rate of

the system. Increased dilution rates result in increased nutrient and decreased waste

product concentrations. Thus, a high degree of control can be exhibited over the

environment of this system One obvious advantage of continuous culture lies in

elimination of the requirement for repealed growth cycles. Continuous cultures are

generally operated well below the maximum growth rate, therefore reducing the medium

wasted on biomass production. Operating expenses are also comparatively lower for

continuous culture as a result of more efficient labor use and reduced turnaround costs.

However, perfusion suffers from some of the same disadvantages as the batch method

with periodic replenished of the medium, that is, nutrients are being discarded in the

effluent of the culture vessel, which is not cost effective.

2.5 Fibroblast Cell Culture

A fibroblast cell has a spindle shape and is found in the connective tissue in association with collagen. Fibroblast cells are widely used in studies of malignant transformation and control of cell growth and differentiation (Goldberg, 1977). The

fibroblast cell line (3T3) was developed from minced and trypsini/ed whole embryos of inbred Swiss (Todaro and Green, 1963), NIH, or inbred Bai7c mouse strains (Aaronson and Todaro, 1968). The original 3T3 cells were judged to be fibroblasts because, first, they synthesized collagen at the same rate of f ibroblasts (Green, Todaro and Goldberg, 17 1966); second, typical collagen fibers formed in the cultures (Goldberg and Green, 1964,

Todaro, Green and Goldberg, 1964); third, they secreted proteoglyccans (Hamerman,

Todaro, and Green, 1965); and fourth, they can assume a fibroblastic morphology after

viral transformation (Todaro el al., 1964).

Fibroblast cells are typical anchorage dependent cell lines and have been used in

many mammalian cell culture studies, because they can survive most mechanical and

enzymatic exploitation techniques, can be used as host cells for recombinant DNA expression, and can be cultured in many simple media, such as Eagle s basal medium.

Ro/.engurt et al. (19X7) used Swiss 3T3 fibroblast cells to study the activation function of

protein kinase C on cAMP accumulation in the cells. Human diploid fibroblasts were cultured and induced by virus and dsRNA to produce interferon, an inducible secretory protein produced by eucaryotic cells in response to viral infections (Friedman, 19X1). By culturing human foreskin fibroblasts in microcarricrs, Croughan et al. (1988) achieved a high cell density (2-3 x 1()6 cells/ml), and also studied effects of microcarrier concentration and hydrodynamic forces (Croughan et al., 19X7). 3T3 fibroblast cells also have been used frequently for fibroblast cell physiological and morphological studies and for gene expression after being transinfection by certain DNA encoding genes (Freshney,

1991). Caldwell et al. (1991) studied medium exchange schedules in genetically engineered NIH 3T3 cells and determined the relationship between the consumption rate of medium components and generation rate of metabolites.

During culturing of fibroblast cells, the following cyclo-oxygenase pathway occurs (Benedetto et al, 19X7): the first step is the formation of the hydroperoxide of the parent straight chain C 20 polyunsaturated fatty acid. After a series of enzymatic transformations, prostaglandin G 2 (P G G 2 ), catalyzed by cyclooxygenaxe, is produced.

A lte r P G G2 is hydro-oxided to form PGH 2 by hydroperxidase, prostaglandin E 2 (PG E 2), a relatively stable intermediate, is generated and secreted with both reductase and IK isomerase enzymes. The cyclo-oxygenase and lipoxygenase pathway is shown in Figure

2.1. Phorbol 12-myristate 13-acctate (PMA), a drug extracted from plants, can stimulate

the expression of cyclooxygenasc via protein kinase C, resulting in over-production and

secretion of prostaglandin E 2 (PG E 2 ) Since the productivity of prostaglandin E 2 is

proportional to the cell density of fibroblast cells, the cell growth rate, product

productivity, and culture system efficiency can be assessed by measuring the

concentration of PGE2 and cell density.

2.6 Primary Luteal Cell Culture

The corpus luteum (CL) is a transient endocrine gland formed within the ovary by

differentiation of follicular cells after ovulation. If an embryo is not present, the corpus

luteum w ill regress, and a new corpus luteum w ill form during the next ovulation. The

primary role of the corpus luteum is to synthesize and secrete progesterone, which is

essential for the maintenance of pregnancy (Smith, 1986). Development, maintenance,

and regression of the corpus luteum have been investigated for many years. However,

endocrine and cellular mechanisms regulating progesterone synthesis and secretion are

still unclear (Smith. 1986).

Corpus luteum can be enucleated from animals, and after dissection, a large

number of luteal cells can be attained. Although luteal cells can synthesize cholesterolde

novo, this cholesterol source is not sufficient to meet the demands of steroidogenesis.

Therefore, luteal cells are largely dependent upon extracellular cholesterol sources, and

large amounts of cholesterol are required to serve as the substrate for maximal

steroidogenic production (Pate and Condon, 1982). Progesterone production is regulated and maintained by luteotropic and lutcolytic mechanisms, or by the pituitary hormone and luteinizing hormone (LH) (Hansel et al., 1978; Rothchild, 1981). The specific roles of luteinizing hormone in progesterone synthesis is currently unclear. A relatively 19 reasonable model proposed by Niswender ei al. (1980) are the second messenger theory.

According to this theory, LH stimulates progesterone synthesis by a second messenger

system, which involves a protein hormone (LH, 1st messenger), membrane-bound

receptor, conversion of adenosine triphosphate (ATP) to adensine-3',5'-cyclic

monophosphate (cAMP; 2nd messenger) by adenylate cyclase, activation of a cAMP-

dependent protein kinase, and an intracellular response. At first, luteinizing hormone

binds to a specific membrane receptor and activates adenylate cyclase resulting in cAMP

synthesis. The hormone receptor complex is internalized subsequently by endocytosis

and degraded within a lysosome. Increased intracellular cAMP concentrations stimulate

protein kinase activity in cell cytoplasm, which may phosphorylate steroidogenic

enzymes (i.e., cholesterol esterase and side-chain cleavage complex), and the produced

cholesterol is transformed in mitochondria to pregnenolone after a series of enzymatic

reactions. Pregnenolone is released from the mitochondria and converted to progesterone

in the smooth endoplasmic reticulum. The final result is the increase of progesterone

synthesis. The metabolism of stimulation of luteinizing hormone and lipoproteins to

luteal cells and the production and secretion of progesterone is shown in Figure 2.2.

The corpus luteum contains a heterogeneous population of cells that have different size, appearance, organelles, and steroidogenic capability (Foley and Greenstein,

1958; Koos and Hansel, 1981). Steroidogenically active luteal cells include both small and large cells (Urseley and Leymarie, 1979; Koos and Hansel, 1981). It was reported by

Alila and Hansel (1984) that small and large luteal cells originated from theca and granulosa cells, respectively, and small luteal cells developed into large luteal cells as the luteal process progressed. There are distinct differences in the ability of small and large luteal cells to secrete progesterone in the presence of LH (Urseley and Leymarie, 1979;

Koos and Hansel, 1981). Specifically, large bovine luteal cells secrete 20 times as much 20 progesterone as small luteal cells. However, small luteal cells are 6 times more responsive to LH than large luteal cells (Urseley and Leymarie, 1979).

The physiological properties of luteal cells in vitro have been studied for a long lime. The effects of serum and lipoproteins on the function of bovine luteal cells were examined by Pate and Condon (1982). It was found that the presence of serum in the cell culture medium inhibits the responsiveness of luteal cells to LH and lipoproteins can be used to increase progesterone production. Pate et al. (1987) also found that the capability of progesterone production by luteal cells decreased the culturing time even with the addition of LH. Many hypotheses have been proposed to explain this phenomenon.

Besides the lower cell density compared with that in vivo (Pate et al., 1987), Chow and

Poo (1982) reported a redistribution of myotomal cell surface lectin receptors when cell- cell contact occurred. Alterations in the arrangement of these receptors could certainly change cell response to the hormone. Receptor-mediated events may also be influenced by cell density via alterations in plasma membrane fluidity (Danforth et al., 1985).

L is well documented that mammalian cells cultured in vitro will either retain their morphological and biochemical differentiation, and, thus, their protein synthetic and secretory capabilities, or will lose them depending on the chemical composition and structure of the culture supporting substrate. It has also been clearly demonstrated by

Emerman and Pitelka (1977) and Ruzicka (1986) that cells cultured on flat plastic surfaces or on two-dimensional collagen sheets more or less lose their differentiated ability for protein synthetic and secretory activities. Thus, it is important to consider two factors in developing an optimal culture condition and materials for mammalian cells to maintain their differentiated properties: (1) The correct material composition, and (2) the proper shape of the substrate (Yang and Nandi, 1983). According to Yang's conclusions, both the shape of a cell and its orientation to neighboring cells were important in modulating the proliferative response of a cell to mitogen. Folkman et al. (1978) and 21 Gospodarowicz ct al. (1978) have shown that cell shape was tightly coupled to DNA and

protein synthesis, protein secretion, and cell growth.

It has been highly desirable to employ a culture system in which the function of

cells is close to that in vivo. Gospodarowicz and Gospodarowicz (1972) reported that

bovine luteal cells were successfully cultured in vitro with static flasks. Stoklosowa and

Sladnicka (1973) successfully grew rat luteal cells in vitro, and Goldsmith and his

colleagues (1981) developed a continuous culture system for human luteal cells.

However, these systems had some problems and were not extensively used to study luteal

steroidogenesis in vitro. Almost all the experiments so far have been done only in static

flasks. The disadvantages of culturing luteal cells in flasks are: (1) poor mass transfer because the diffusion is the only mass transfer pattern in the static culture, (2) nutrient depletion and accumulation of toxic metabolites which are deleterious to cells, at the end of the culture, and (3). alteration of cell morphology, because the surface of static flasks can provide only a two-dimensional growth environment for cells to grow.

2.7 Hybridoma Cell Culture

Kohler and Milstein (1975) demonstrated that individual clones of normal antibody secreting cells could be immortalized by fusion with myeloma cells. Fusion resulted in formation of hybridoma cell lines, which secreted the antibodies of the antibody forming cells and could be propagated indefinitely in vitro like the myeloma cells.

Mammalian cells have two pathways for DNA biosynthesis: the de novo synthesis pathway and the salvage pathway. The de novo pathway is the main biosynthesis pathway for purines and pyrimidines and can be blocked by the folic acid antagonist, aminopterin. Aminopterin-blocked cells can still synthesize DNA via the salvage pathways in which preformed nucleotides are recycled. The salvage pathways depend on 22 the enzymes thymidine kinase (TK ) and hypoxanthine guanine phosphoribosyl

transferase (HGPRT). Thus, if the aminopterin-blocked cell is provided with thymidine

and hypoxanthine, DNA synthesis can still occur provided the enzymes TK and HGPRT

arc present.

If either TK or HGPRT enzymes are absent , aminopterin-blocked cells can not

grow because they cannot synthesize DNA. However, the blocked cell can be rescued by

fusion with another cell which supplies the missing enzyme. This process is invented by

Bulter (1987) and now is commonly used in hybridoma selection protocols. A myeloma

cell lacking TK or HGPRT is fused with a normal spleen cell that possess the missing

enzyme . The fused hybrid is cultured in medium containing aminopterin, hypoxanthine

and thymidine (H A T medium) and only the hybrids (myeloma x spleen cell) live. The schematic diagram of the production of hybridoma cells secreting monoclonal antibodies is shown in Figure 2.3.

2.7.1 Monoclonal Antibody Production

The techniques of monoclonal antibody production by culturing hybridoma cells have been studied for several decades and research have made great progress in this area.

Compared to in vivo incubation, cell culture has a number of distinct advantages: (1) it has potential for unit operation, therefore for reduction of unit operation cost as scale increases, (2) it reduces risk of contamination of the product by rodent-derived infectious agents, (3) it avoids the presence of extraneous rodent antibody, as found in ascites, and

(4) the process can be engineered and controlled to be highly reproducible.

2.7.1.1 Suspension Culture

Suspension cell systems are frequently run on a batch basis, in which the cells produce antibody for only a few days before they die. This causes difficulties in down 23 stream processing, requiring for removal of dead cells from the antibody-containing medium. The batch systems achieve a relative low concentration of biomass (magnitude of 106 cells/ml in batch system and 107 cells/ml in perfused suspension) compared to ascites cultures (108 cells/ml) or to the growth of animal cell in tissues ( 109 cells/ml)

(Altshuler et al., 1987; Katinger, 1987). Daliliand Ollis tl990) cultured hybridoma cell line MRC O X -19 in 100 ml T-flasks for 170 hours and the highest cell density and monoclonal antibody was about 2xl()6 (cell/ml) and 21 (mg/L), respectively. However, since batch production is a well-understood technique and easy to scale up, most large industrial productions still use batch reactors. Backer and his colleagues (1988) at Eli

Lilly & Company reported that by using a marine impeller agitator they successfully cultured hybridoma cells in a 1300-liter fermenter. Lambert et al. (1987) grew 35 different monoclonal antibody lines in 1000 liter airlift system, and antibody production levels ranged from 40 to 500 mg/L.

Many modifications have been made to increase the cell density and productivity of the suspension bioreactors. Broise et al. (1992) coupled a stirred tank to external tangential flow filtration, and the cell density and MAb concentration increased to SxlO6

(cell/ml) and 120 (mg/L), respectively. By using an air-lift bioreactor connected by a cell settler for cell retention, Hulscher and his colleagues (1992) increased the MAb productivity by a factor of 17 and the cell density by 4 compared with conventional batch systems. Shi et al. (1992) designed a new impeller to improve the oxygen transfer in the bioreactor, and the total hybridoma cell concentration was increased to 3.4 xl()7 cells/ml, and the MAb concentration was also increased to 512 mg/L. Increase of cell density and

MAb production by using a dialyzed continuous suspension culture system was also reported by Linardos et al. (1992). 24 2.7.1.2 Immobilized Cell Culture

Most hybridoma cells can be cultured in conventional fermenters. However, it is

more desirable to culture the cells under immobilized conditions, because the hybridomas

require a stable, shear-stress-free physical and chemical environment to achieve optimal

growth and productivity (Emery et al., 1987).

Immobilization techniques that seek to make use of a more or less conventional reactor design involve in entrapping the cell within some form of particles, such as agarose (Nilsson, et al., 1983), alginate (Siacore, 1984), or a membrane-bound capsule

(Rupp, 1985). The immobilized cells are maintained in a protected, biochemically conditioned environment. Beads and porous particles allow a cell-free product stream to be withdrawn from the reactor. In all of these cases, the reactors can easily be perfused with a continuous supply of fresh culture medium, extending the productive lifetime of the cells and increasing the cell concentrations obtained in the reactors.

Novel bioreactor designs that do not use particle-entrapped cells achieve immobilization by enclosing the cells within one or more specific compartments.

Examples include many variants of hollow-fiber bioreactor (Altshuler, et al., 1987;

Tharakan et al., 1986), in which cells are enclosed within the shell of a hollow-fiber cartridge, and the membrane bioreactor design of Klement et al.(1987), in which cell are contained in a thin layer between two planar membranes. These systems, too, maintain the cells in a conditioned and protected environment. A cell-free and, in some cases, concentrated product stream can be withdrawn from the reactors, and cells can be perfused with fresh medium to increase cell concentration and the life time of the culture.

Lee (1991) compared monoclonal antibody production in suspension cell and immobilized cell culture (S3H5/y2bA5 hybridoma cells in alginate beads) and the experimental results showed that the viable cell density and the volumetric MAb productivity were 3 time as high as those in the suspension cell culture. Cadic and his 25 colleagues (1992) immobilized hybridoma cells in agarose beads and achieved 20-fold

more cell density than that in suspension culture. Broise et al. (1992) cultured the

hybridoma cells in both alginate membrane and hollow fiber bioreactor and achieved 1.5-

2 folds increase of MAb production, respectively. A better result was obtained by

Altshuler (1986) and his colleagues (4.5 and 7 folds, respectively) using a hollow fiber

bioreactor. Hagenom and Kargi (1990) developed a coiled tube bioreactor and achieved a 6 folds increase of MAb increase compared with spinner flasks. New Brunswick

Scientific Co., Inc. (Wang et al., 1992; Biomass News, 1993) invented a packed-bed

bioreactor (CelliGen) in which the productivity could reach 12-fold high of that in static and stirred suspension culture systems. A comparison of suspension cell and immobilized cell culture is shown in Table 2.3.

Inevitably, such improvements in performance also raise some problems. Cadic and Dupuy (1992) reported that the production of M Ab decreased during two weeks cultivation because of the spherical colony formation, which caused mass transfer limitation. Broise et al. (1992) also observed the production of MAb declined after 3 weeks cultivation because of the destruction of the beads. The fiber-bed bioreactors can achieve a high cell density, however, it suffers clogging and channeling problems with time. Experiments by Wang et al. (1992) showed that the productivity of MAb in the packed-bed bioreactor (CelliGen) started decreasing after 35 days cultivation. The reasons behind these phenomena were complicated. However, poor mass transfer played an important role, especially after a large amount of cell mass built up inside the bioreactors.

2.8 Hybridoma Cell Kinetics and Simulation of the Fibrous Bed Bioreactor

Numerous models have been developed to describe the growth of hybridoma cells, production of monoclonal antibodies, consumption of nutrients in the medium, and 26 generation of metabolites. Hybridoma cells, like other mammalian cells, grown in vitro

exists in a dynamically complex and ill-defined environment (Glacken et al., 1988).

Thus, understanding the roles of physiological and environmental factors on the growth

and metabolism of hybridoma cells is a prerequisite for the development of rational scale-

up procedures (Miller, 1988).

Although there been much literature published addressing hybridoma growth and

MAb production, most of them were based on experimental data from static flasks and suspension culture. Bibila and Flickinger (1991, 1992) proposed a structured kinetic model to describe the correlation between cell growth phase and MAb synthesis and used this models to optimize the process of MAb production. Based on their results, environmental and/or genetic manipulation approaches could maximize the specific antibodies secretion rate and the antibody volumetric productivity in large-scale antibody production systems. M iller et al. (1988) and Ozturk et al. (1991) independently conducted a kinetic analysis of the effects of nutrient, pH, and dilution rale on hybridoma growth and metabolism in batch and continuous suspension culture. A schematic diagram of the metabolic pathway of hybridoma cells proposed by Miller et al. (1988) is shown in Figure 2.4. The rheological properties of hybridoma cells were studied by Shi et al. (1993). Dalili and Ollis (1989), Heath (1989), and Lee (1991) investigated the effects of serum and other energy sources on cell growth, metabolism, and antibody production. The kinetic analysis by Glacken et al. (1988, 1989) suggested that the growth rate of hybridoma cells could be described by the Monod equation, glutamine had a noncompetitive inhibition effect on cells, and lactate was the only environmental parameter that significantly inhibited antifibronectin MAb production by CRL-1606 hybridomas. Bree et al. (1988) studied the relationship of glucose, glutamine, and lactate and suggested the rate of glucose consumption was zero order in glucose concentration rather than following the Monod model. Ozturk et al. (1988) and Miller et al. (1988) 27 concluded that the production of monoclonal antibodies was non-growth associated.

However, in batch culture the period over which the specific growth rate varies is small, and it is difficult to observe the range of specific growth rates possible with continuous culture (Miller and Blanch, 1991). Glacken et al (1988) showed that the specific rate of

MAb production was growth associated up to a specific growth rate of 0.02/hr and independent of Specific growth rate (|i) at growth rate greater than 0.02/hour and the specific production rate also depended on lactate concentration.

Mathematical modeling of mass transfer phenomena and cell kinetics can indicate the relationship among nutrient, metabolite, and product concentration inside the bioreactors (Heath and Belfort, 1991). Many papers describing theoretical models of substrate diffusion/convection and uptake in different bioreactors have been published.

Park and Chang (1986) described the flow distribution in a hollow fiber bioreactor and the flow patterns were also visualized and modeled with magnetic resonance image techniques by Heath et al. (1990). Models were developed for packed-bed and fibrous bed bioreactors. Ethier (1981) investigated creeping flow through mixed fibrous porous materials and found that vicious effects at the coarse fiber surfaces led to a significantly lower overall permeability than that predicted by a simple application of Darcy's law.

The pressure drop in a fixed bed of spherical particles was measured, calculated, and predicted by Jaiswal et al. (1994). Avlontis (1993) and Boundinar (1992) improved the performance of spiral-wound (SPW) modules by optimizing some key geometrical parameters for given operating conditions. When the packed bed and fibrous bed bioreactors were employed to cultivate microorganisms and mammalian cells, many models coupling mass transfer equations with bioreaction kinetics were developed to simulate and optimize the bioreactors. Perry and Wang (1989) addressed the fundamental requirements of fibrous bed bioreactor for mammalian cell culture. Chiou et al. (1991) and Murakami (1991) modeled a concentric-cylinder airlift, fibrous bed reactor 28 for mammalian cell culture and also studied its scale-up potential. The study of Park and

Stephanopoulos (1993) on a ceramic bead packed-bed cell culture bioreactor found that the extent of intraparticle convective medium flow was the dominant mechanism of nutrient transport to cell immobilized in the particle interior and the model thus developed was helpful for enhancing the limiting nutrients, such as oxygen, to allow maintenance of cell viability and productivity. The mass transfer and bioreaction kinetics in the biofilm on the surface of the packing material in the packed bed bioreactor was modeled by Skowlund and Kirmse (1989) and Lewandowski et al. (1991). Interactive dynamic programming (IDP) technique was used by Hartig and Keil (1993) for optimization of large scale, spherical, fixed bed bioreactors and was found the global optimum with higher probability compared with conventional methods. Despite their limitations, computer simulations can provide a significant savings of time, effort, and expense by allowing a quick and easy generation and inspection of bioreactor concentration profile. Design, operation, and scale-up may also be significantly improved with the insights offered by the flexibility of theoretical investigations. 29 Table 2.1. Classification of Mammalian Cell Bioreactor According to Reactor Type (Prokop and Rosenberg, 1989)

Reactor type Type of support substrate Suspension (S) Anchored (A)

Stirred tank reactor Possible for all: A and S Rat turbine Microcarriers A and S Stirrer bar/paddle Microcarriers A and S Paddle w.fiat sheet Porous support A and S Angled blades Beads A and S Marine propeller A and S Vibromixer not recommended A and S for microcarriers Plough-shaped sheet A and S Helical mixer A and S Anchor mixer A and S Helical w.draft tube A and S Ribon and screw A and S Cavity (cage)type A Hanging stirrer A and S Roat type A and S Horizontal loop type A and S Film type A and S

Column reactor Column Rotating vertical stack plate A Rotating horizontal disc A Bubble column Porous support or beads A and S Column w. draft tube Porous support or beads A and S Ruidized bed Porous support or beads A ( simple column) or staged) Packed bed Glass spheres A Ceramic or glass tubes A

Membrane reactor Hollow Fiber Capillary (extracapillary A and S or lumen) Rat membrane Dialysis reactor No support S Dialysis bag/fleaker No support S Maintenance reactor No support S (membranes for Porous support A nutrient supply) Microcarriers A Ceramic monolith Ceramic support A and S Microencapsulation Capsule A and S Gel matrix Gel bead A and S 30

Table 2.2 Surface -volume ratio (cm-1) of various methods to culture anchorage- dependent cells (Glacken et al., 1983)

Type of bioreactor S/V ratio (cm*1)

Plastic bags 50 Multiple propagator 17 Spiral Film 40 Glass bead propagator 100 Artificial capillaries 307 Microcarrier suspension 122-153 Jensen's IL410 tubular spiral film 94 Gyrogen with tube 12 ♦Fibrous Bed 96.5

* The fibrous bed bioreactor used in this research Table 2.3. Comparison of Suspension and Immobilized Cell Culture Techniques

Technique Hybridoma Antibody Cell Density Yield Concentration Reactor Volume References (log cells/ml) (mg Ig/day) (mg/L) (mL)

Free Cett Suspe/inoH

Batch Ascties mouse IgG 7.3 30x1 O'3 5000 4 (per mouse) Suspension mouse IgG 4.0 027 5-20 100 Altshuler, et al., (1986) Culture Rat IgG 6.3 10.2 5 - 30,000 Birch et al., (1984)

Chemostat mouse IgG 5.3-6.0 0.92 - 43.3 2-80 500 Fazelas., (1983)

Immobilized cells

Gel rat IgG 7.3 (ested) 200 5 Cadic et al., (1992) mouse y2b 626 25-30 20 Lee et al., (1992)

Ceramic mouse Ig 55 Putman et al., (1985)

Hollow mouse IgG 6.5 -7.3 4 400-750 (shell) 2.5 Altshuler, et al., (1986) fiber 260 (reservoir) 250 Altshuler, et al., (1986)

Flat mouse IgG 8-8.4 28.5 - 570 12.5 (cell) Seaver et al., (1985) membrane

spherical mouse IgG 8 3.13 2500 (capsule) 12.5 Littlefield et al.,(1984) membrane

Packed mouse IgG 8 500 215 1800 Wanget al., (1992) Bed 32

Arachidonatc

('wbvi^irnu1^ k » . - COOH CH COOH I*GG. l« II 14 ^•ciM idnm oiie OOH

r 2/3 '-'♦H .O

O PGII, COOII

O CH OH

CH COOH CH ,CH,

OH OH OK

Thfomboiune A \ OH

COOH COOH

COOH

Figure 2.1 The cyclo-oxygenase and lipoxygenase pathway of arachidonic acid metabolism (Rawn, 1989) 33

MEMBRANE CYTOPLASM MITOCHONDRIA

\ AA« PROTEIN CHOLESTEROL AC Imoctivt) CJ—*|RJ |C I POLYPHOSPHOINOSITIDE

CHOLESTEROL LIPIO DROPLET h lAOPH LH- ? / CHOLESTEROL ' ESTER CHOLESTEROL P-LIPIDS CHOLESTEROL ♦ TEA 4-ZCHj +CMj f PREGNENOLONE \ i * : ^PROGESTERONE^

►cholesterol ► CHOLESTEROL

ACETATE ♦♦CHOLESTEROL

Figure 2.2 The metabolism of stimulation of luteinizing hormone and lipoproteins to luteal cells and the production and secretion of progesterone Figure 2.3 Schematic diagram of the production of hybridoma cells secreting monoclonal antibodies (Nisonoff, 1984) 35

OIm m m

N A O H ♦ W * .

QLYCOLY5IS Qtyctnu

Amino Acid*

NAOH n Pynm i»

/ MO

A tp a n w M

TCA CYCLE

NAOH ♦ O

AOP

Figure 2.4 Schematic diagram of hybridoma cell metabolic pathways ( Miller et al, 1988) CHAPTERID

FIBROBLAST CELL CULTURE

3.1 Introduction

The key of designing a high-efficient bioreactor for culturing anchorage

dependent cells is to increase the ratio of the surface area available for cell growth to the

total bioreactor volume. Various studies have been done on the design of novel materials

and bioreactors to meet this need. Van Wezel 31 developed the concept of microcaniers,

which offers a large surface per unit culture volume for cell attachment and grow cells

from an initial concentration of 2 x 10s cell/ml to 2 xlO 6 cells/ml. The microcaniers now

available from commercial suppliers (e.g. Flow's Superbeads and Pharmacia's Cytodex)

can achieve much higher cell densities. However, according to Griffiths ,16 and Robert et

al., 26 scale-up of microcarrier bioreactors remains highly problematic. The biggest limitations are the development of suitable inoculation protocols ,5-13 the culture's extreme sensitivity to mixing and bubble shear, 5-25 and the high cost of beads . 12 The potential of using hollow fiber filtration devices for high density cell growth was first recognized by Knazek et al .21 Cells are inoculated on the shell side while medium and air flow through the lumen of each fiber. The cells are retained in the capillary space and grow to a high density. Hollow fiber bioreactors have been successfully used for production of monoclonal antibodies , 2 hormones ,22 cancer cells , 18 and viral antigens . 19-23 By culturing hybridoma cells in a hollow fiber bioreactor, cell density can be increased by two orders of magnitude and the productivity can be increased by 20 to

40 times.2®

36 37 A fixed bed packing matrix has to be a compromise between maximizing the

surface area and providing a passage open enough for medium to be perfused at a rate

sufficient to maintain homogeneity throughout the bed without subjecting cells to

damaging shear effects . 16 Many different packing materials have been used with Fixed

bed bioreactors suitable for animal cell cultures .28 A fixed bed mass culturing system

was developed by Brown et al .,8 by which a variety of cells were successfully cultured.

About 1.3 x 10 10 viable cells were recovered from a 12 liters bioreactor, and the system

was kept continuously running for over 1 year. Chiou and his colleagues 10 developed a

cell culture system which combined the features of air-lift and a packed bed of glass

fibers and achieved a high density of CHO cells (1.2 x 10 7 to 5.1 x 10 7 cells/ml). New

Brunswick Scientific Co., Inc. 3*31 developed a packed-bed bioreactor (CelliGen) with productivity as high as 12-fold of that in static and stirred suspension culture systems.

Inevitably, such improvements in performance also raise some problems. Cadic and Dupuy 9 reported that the production of MAb decreased significantly during two weeks cultivation because of the formation of spherical colonies causing mass transfer limitation. Broise et al. 7 also observed the production of MAb started to decrease after 3 weeks cultivation due to the destruction of the beads. The fiber-bed bioreactors 10 may achieve a high cell density, however, it suffered clogging and channeling problems after a period of time of operation. Experiments by Wang et al .32 showed that the production of

MAb in the packed-bed bioreactor (CelliGen) started to decrease after 15-35 days cultivation. The reasons behind these phenomena are complicated. However, the poor mass transfer plays an important role, especially after a large amount of cell mass built up in the bioreactor.

A new continuous, fibrous-bed, immobilized cell bioreactor was developed for continuous cultivation of bacteria to produce organic acids at a high efficiency.34-35*36.

The fibrous bed bioreactor contains a spiral wound fibrous sheet for cell immobilization, 38 with significant spaces between wound layers to allow good mass transfer of nutrients and products. The fibrous bed bioreactor provides a large surface area for cell attachment, which results in a high cell density and productivity. It also provides a conducive three dimensional matrix for cell attachment, allowing cells to maintain their native morphology and differentiated properties. The structured fibrous bed also allows for good multi-phase flows and provides renewable surface for cell immobilization. It has been shown that cell immobilization in this bioreactor was not irreversible as constant growth of new cells and sloughing off dead cells occurred continuously within the reactor. The bioreactor is thus self-renewing and eventually establishes a dynamic steady-state cell population. This bioreactor offers long-term stability for continuous operation without suffering aging, fouling, and degeneration problems normally occurring to conventional packed bed and membrane bioreactors. This novel fibrous-bed bioreactor is ideal for mammalian cell culture and was used in this study.

The fibroblast cell lines were originally developed from minced and trypsinized whole embryos of inbred Swiss , 30 NIH, or inbred Bai/c mouse strains .1 They are anchorage-dependent cells and have been used in many mammalian cell culture studies or as host cells to produce recombinant proteins as they can survive most mechanical and enzymatic expression techniques. Fibroblast cells have been used frequently for cell physiological and morphological studies and gene expression after transinfected by certain DNA encoding genes .13 Fibroblast cells are widely used in studying malignant transformation and control of cell growth and differentiation . 15 Rozengurt et al. 27 used

Swiss 3T3 fibroblast cells to study the activation function of protein kinase C on cAMP accumulation in the cells. The functions of protein kinase C in NIH 3T3 cells on cell growth, morphology, anchorage dependence, and tumorigenicity were studied by

Mischak.24 Croughan et al .11 achieved a cell density of 2~3xl0 6 cells/ml by culturing 39 human foreskin fibroblasts in microcarriers and also studied effects of roicrocarrier

concentration and hydrodynamic forces .12

Lipoxygenase pathway occurs in the metabolism of fibroblast cells , 4 during which

prostaglandin E 2 (PGE2) catalyzed by cyclooxygenase is produced and secreted. Phorbol

12-myristate 13-acetate (PMA), a drug extracted from plants, can stimulate the

expression of cyclooxygenase via activating protein kinase C to result in over-production

and secretion of prostaglandin E 2 (PGE2). By measuring the concentration of PGE 2, cell

growth conditions, cell density, and the productivity of fibroblast cells in different

systems can be assessed.

In this study, fibroblast cell NIH 3T3 from mouse were cultured in static T-flasks,

microcarrier bioreactors, and a fibrous bed bioreactor, respectively. The production of

FGE2 as stimulated by PMA in various culture systems was studied and used as a

measure to evaluate the efficiency of different culture systems.

3.2 Materials and Methods

3.2.1 Cell Line

Fibroblast NIH 3T3 cells were obtained from Dr. Douglas Kniss, Department of

Obstetrics, The Ohio State University Hospital.

3.2.2 Medium

The medium composition was as follows: 90% (v/v) DMEM (4g/l glucose, Difco

#33859-017), 10% (v/v) new bom calf serum (Difco #16010-019), 0.29 g/1 glutamine

(Difco #21051-024), and 0.06 g/1 gentamicin (Whittaker #1301-898-7025). The medium pH was 7.2. The medium was sterilized by filtration through a 0.22-Jim medium filter

(Coming #25970-33). 40 3.2.3 Determination of Optimal PMA Concentration

A total of 5xl0 5 fibroblast cells were transferred into each of four T-175 culture flasks, into which fresh medium was added to make the final liquid volume of 2 00 ml.

After 72 hours cultivation, cells became confluent in the flasks. Then, the medium in the flask was changed by adding fresh medium with different concentrations of PMA. Four different levels of stimulus (PMA) concentration (10-6, 10'7, 10*8, and 5xl0 ' 9 M) were tested in different flasks. The flasks were then incubated in a CO 2 incubator (Napco E series, Model 302) at 37 °C. During the culturing period, samples were taken at different time points and prostaglandin E 2 was measured by radioimmunoassay. By comparing the production of prostaglandin E 2, the optimal PMA concentration was determined and used for the rest of the experiments.

3.2.4 Culture in T-Flasks

T-175 flasks were used for static culture of fibroblast cells. About l.OxlO 8 cells were transferred into each T-flask, and fresh medium was added to a final total volume of

200 ml. The T-flask culture was incubated in a CO 2 incubator at 37 °C. During the incubation period, 12 ml medium were taken out of the flask and the same volume of a fresh medium containing optimal PMA (10 *7 M) concentration was added into the flask every 4 hours to make the static culture close to the condition of continuous culture. The removed medium samples were frozen for future analysis.

3.2.5 Culture in Microcarrier Bioreactors

The procedure of culturing cells on microcarriers was based on the instructions from Pharmacia LKB Biotechnology Co.. Cytodex-2 microcaniers (Pharmacia) were used to grow the fibroblast cells. The physical characteristics of Cytodex-2 are as follows: density: 1.04 g/m l, size: 170 fim, approx. area: 3,300 cm2/g dry weight, and 41 approx. no. microcaniers: 4.1xl06/g dry weight. A total amount of 0.8 grams of

Cytodex-2 was washed and hydrated in 40-80 ml of Ca2+ and Mg2+ free Dulbecco's phosphate buffer saline (PBS) for at least 3 hours. The supernatant was decanted and the microcarriers were washed for a few minutes in fresh PBS again. Then, the washed microcaniers were put in a 500 ml spinner flask (Gibco) and sterilized by autoclaving

(121 °C, 15 psi, 15 min). After cooling down to room temperature, a 200-ml fresh and sterilized culture medium and cells (to a total cell density of 5x10s cells/ml) were transferred into the spinner flask. The seeded flask was incubated in a CO 2 incubator

(Napco E series, Model 302) at 37 °C. The stirring speed for the spinner flask was about

70 rpm. For every 4 hours, 12 ml of medium were removed from the flask and the same volume of a new medium containing 10 ' 7 M PMA was added into the flask. The removed medium samples were frozen for future analysis.

3.2.6 Culture in a Fibrous-Bed Bioreactor

3.2.6.1 Selection of Packing Materials in the Bioreactor

Five different fibrous sheet materials were used to test the attachment condition of fibroblast cells. They were: (1). 100% polyester, ( 2 ). 100% cotton, (3). 50% polyester/50% cotton, (4). 100% glass fiber, and (5). 50% polyester/50% rayon. The fibrous materials were cut into small pieces (lxl cm) and placed in a 12-well culture plate. Cells with final concentration of lxlO 4 cell/ml were inoculated into each well and the medium was changed every 24 hours. After 2 day culturing, the fibrous materials were taken out and dehydrated for examination using scanning electron microscopy

(SEM) to evaluate the suitability of each fibrous material for cell attachment. 42 3.2.6.2 Well-Mixed Bioreactor

The immobilized cell bioreactor was made of a glass column with 4.5 cm internal diameter and 15 cm height. Water of constant temperature (37°C) was circulated through the water jacket around the bioreactor to keep the temperature in the reactor constant. A piece of 50% polyester/50% cotton fiber (100 x 10 x 0.1 cm) was spirally wound and put into the bioreactor on top of 1" thick 1/4" glass beads, which were used to help to form an even distribution of the feed medium. The height of the fibrous packing was 10 cm and the total liquid volume in the bioreactor was about 200 ml. Two peristaltic pumps

(Masterflex, Cat.#N-07520-14) were used in the system. One pumped the fresh medium into the bioreactor, and the second one recirculated the medium in the bioreactor through the recirculation flask at a high flow rate (80 ml/min) to provide well-mixed conditions in the reactor. A pH probe (Cole Parmer, #G-5662-10), a dissolved oxygen probe (Cole

Parmer, #G-5644-00), an inoculation port, and a sampling port were installed on the top of the reactor. A 2-liter spinner flask (with 100 ml liquid volume) was installed in the recirculation loop. Sterilized air containing 5% carbon dioxide were used to flush the surface of the liquid in the recirculation flask to supply the oxygen to the medium and balance the pH in the bioreactor. The effluent overflow out of the spinner flask by the pressure built up inside the system. Figure 3.1 and Figure 3.2 show the schematic diagrams of the spirally wound structure of fibrous matrix and the well-mixed bioreactor system used in this work, respectively.

3.2.6.3 Plug-Row Bioreactor

The well-mixed bioreactor could be converted to a single-pass bioreactor with modifications: ( 1). disconnecting the recirculation loop and converting the spinner flask to a feeding tank, in which the fresh medium was surface-aerated (95% air and 5% CO 2) before being pumped into the bioreactor, and ( 2 ). moving the effluent port to the top of 43 the bioreactor. A schematic diagram of the plug-flow type bioreactor system is shown in

Figure 3.3. In this single pass bioreactor, concentration and pH gradients existed along

the reactor length. Since the feeding stream passing through the glass bead layer had the same concentration at all radial position and the flow velocity along the bioreactor length was small (1 cm/hr), the flow condition in the bioreactor was close to plug flow.

3.2.7 Bioreactor Start-up

The bioreactor system, including all feed tanks, flasks, and tubing connections were autoclaved at 121°C for 1 hour at least twice with a 24-hour interval. One liter of filter-sterilized medium was prepared in a 2-liter flask. After the bioreactor cooled down to the room temperature, the medium was pumped into the reactor until the reactor was full. The water bath connected to the water jacket of the reactor was started working to keep the temperature of the bioreactor at 37°C. Concentrated cells were injected from the inoculation port using a 10 ml syringe to a total cell density of 5 x 10 5 cells/ml in the bioreactor. The recirculation was turned on at a slow speed (20 ml/min) to ensure that the cells were evenly distributed in the fibrous matrix. No fresh medium was added into the bioreactor until the pH dropped to 6.8 or dissolved oxygen to 15%. Then, a low pump speed (0.2 ml/min) was set to replenish the reactor. The pump speed was increased gradually until it reached 0.8 ml/min. It took about 35 days for the bioreactor to reach steady-state as determined from the stable outlet glucose and lactate concentrations. The fibrous bed bioreactor was then used in kinetic studies under batch, continuous well- mixed, and plug-flow conditions, in sequence.

3.2.8 PGE 2 Production in the Bioreactor under Batch Condition

After the fibrous bed bioreactor was saturated with fibroblast cells, the first experiment was carried out at batch mode. In this batch experiment, the medium in the 44

bioreactor was quickly changed with fresh medium containing 10 *7 M PMA and the feed

,pump was then stopped for ~65 hours while the recirculation pump was still on. Liquid

samples were taken from the bioreactor at various time points and frozen for future analysis of glucose, lactate, glutamine, and PGE 2 concentrations. The experiment was continued for -3 days, and the bioreactor was then changed back to continuous operation under well-mixed condition to restore steady-state for further kinetic studies.

3.2.9 PGE 2 Production in the Bioreactor under Continuous Conditions

The continuous bioreactor was operated at well-mixed mode for about 50 days and then at plug-flow mode for about 50 days. Fresh medium containing 10 *7 M PMA was fed into the bioreactor at various flow rates controlled by pump speed. Effluent samples were taken from the bioreactor outlet and frozen for future analysis of the concentrations of prostaglandin E 2 (PGE2), glucose, lactate, and glutamine. For each retention time studied, pseudo-steady state samples were taken after at least two bioreactor volume of effluent had been collected. After each retention time experiment, the feed rate to the bioreactor was changed to 0.17 ml/min (retention time: 20 hrs) to restore the bioreactor to its optimal condition. The recovery time was dependent on the experimental conditions.

3.2.10 Analytical Methods

3.2.10.1 Cell Density in Static Flasks

Cells were removed from the surface of T-flasks, with 0.5% trypsin in 0.1%

Dulbecco's phosphate buffer saline (PBS), and were then counted on a hemocytometer. 45 3.2.10.2 Glucose and Lactate Concentrations

The concentrations of glucose and lactate in samples were measured using high

performance liquid chromatography (HPLC). The HPLC system consisted of a high

pressure pump (Waters, Model 6000A), an injector (Waters, U 6 K), an organic acid analysis column (Bio-Rad, Model HPX-87H; condition: 45°C), a column heater (Bio-

Rad, a RI detector (Waters, Series 410 differential refractometer; conditions: scale factor

20, sensitivity 16, and internal oven temperature of 45°C), and an integrator (Spectra

Physics, Model 4270; conditions: chart speed 0.25 cm/min, attenuation 32; minimum area

5000, peak height 1, and run time 20 minutes). The effluent was 0.01 N H 2SO4 at 0.6 ml/min flow rate. The details can be found elsewhere35.

3.2.10.3 Glutamine Concentration

The glutamine concentration was determined using an enzyme assay kit from

Boehringer Mannhein (Cat. No. 139092), following the manufacturer's instruction. The detailed description of the experimental procedure is given in Appendix C.

3.2.10.4 Prostaglandin E 2 (PGE2) Concentration

The concentration of prostaglandin E 2 (PGE2 ) was analyzed using radioimmunoassay (RIA ).29 The general procedure is as follows:

100 pi of prostaglandin standard was diluted with Tris buffer to a concentration of

1.56-10 ng/100 pi, and the samples were prepared in 10x75 mm glass tubes in duplicate.

100 pi of antibody to prostaglandin E 2 from chicken dilution (prepared in Dr. Douglas

Kniss's lab, university hospital, OSU) was added to each assay tube (standard curve and samples), and also 100 pi of appropriate dilution of prostaglandin E 2 labeled by 3KTxB 2 was added to each assay tube. After votexed, the tubes were incubated in radioactive refrigerator overnight. 0.5 ml Charcoal/Dextran solution was added to each assay tube 46 and the tubes were votexed. After being incubated again in the refrigerator for 10-15

min., the tubes were centrifuged for 30 min. at 2500 rpm. The supernatant was carefully

poured into scintillation vials and 2.4 ml Scintiverse II (Sigma) was added to each vial to amplify the readings. The assay vials were read in a Beckman (LS 3801) liquid scintillation system, with a multi-sample multi-channel spectrometer. The amount of prostaglandin E 2 in each sample was determined by comparing with the standard curve.

This assay was done in Dr. Douglas Kniss lab in the Department of Obstetrics, The Ohio

State University Hospital.

3.2.10.5 Cell Density in Microcarrier Bioreactors

The cells on microcarriers were harvested by standard procedures (Pharmacia instruction) with proteolytic enzymes, (e.g., trypsin, collagenase).

When harvesting with trypsin, the microcarriers were allowed to settle, the culture medium was removed and the microcaniers were washed for 5 min in PBS containing

0.02% (w/v) ethylenediamine-tetraacetic acid (EDTA), pH 7.6. The EDTA-PBS was removed and replaced by trypsin-EDTA (30-50 ml/g Cytodex). The microcaniers were then mixed well in the trypsin-EDTA and incubated at 37°C with occasional agitation.

After 15 min the action of the trypsin was stopped by adding culture medium containing serum (20-30 ml medium/g Cytodex). The detached cells were then separated from the microcaniers by sedimentation at unit gravity. A certain amount of supernatant was taken out and cells were counted on a hemocytometer.

3.2.10.6 Cell Density in Fibrous Bed Bioreactors

The cell density in the fibrous bed bioreactor was determined from the total protein content in the sample. In general, the total protein content is proportional to the cell number or density. The total protein content was measured by dye-binding protein 47 assay (Bradford method)6. Bovine serum albumin (BSA) was used as the standard for

the protein assay. Five small pieces of fibrous matrix (2.0x2.0 cm) were cut from the

bioreactor packing matrix and used for protein assay and cell density estimation.

Samples and a series of standard BSA solutions (0-100|il, 0.1 mg/ml) were prepared and

mixed with appropriate amount of NaOH (IN) to a final volume of 200 p.1. Then, 5 ml

Coomassie Blue G-250 (Sigma) were added into the tubes and the tubes were votexed.

The standards and samples were read in a Beckman DU 640 spectrophotometer. The

standard curve of protein assay of known cell density is shown in Figure D. 1 in Appendix

D. The cell density present in each piece of matrix sample was then determined from the

protein concentration compared with the standard curve.

3.2.10.7 Scanning Electron Microscopy (SEM)

After all experiments were finished, samples of fibrous matrix (lxl cm) were taken from different positions of the packing. After leaving the samples in 2.5% formaldehyde solution overnight, they were rinsed 10 times by double distilled water, each time for 15 minutes. Then dehydrated with 20 to 70% ethanol in the increment of

10% by leaving them in different concentration solutions for 30 minutes each time. Next day, they were dehydrated with 80% ethanol and then twice with 95% and 100% ethanol for 30 minutes each time. The samples were kept in 100% ethanol overnight again and then dried at critical point with liquid CO 2 in a Pelco critical point drier in the Anatomy

Department of the Ohio State University. All steps, except that for the critical drying, were carried out at 4°C. The dried samples were sputter coated with 60% gold/40% palladium before SEM photos were taken by using JOEL-Model 820 SEM at the Geology

Department of the Ohio State University. 48 3 3 Results and Discussion

3.3.1 Optimal Concentration of PMA

The production of prostaglandin E 2 (PGE2) by cultures stimulated with four

different PMA concentrations (5.0xl0'9, l.OxlO* 8, 1.0x1 O'7, and l.OxlO *6 M) are shown

in Figure 3.4. In general, PGE 2 production was faster with a larger PMA concentration.

However, at l.OxlO ' 6 M, the high concentration PMA was toxic to cells and the

concentration of PGE 2 decreased after reaching a peak level at 5 hours. Increase of

granules inside the cells observed under the microscope suggested that the high PMA

concentration had speeded up the cell death rate. The concentration l.OxlO ' 7 M not only

gave a faster PGE 2 production than did the two lower concentrations, it also kept a

consistent effect on cell growth and PGE 2 production for a long period of time. Also, the

average PGE 2 production was about or more than 10% higher than that from other PMA

concentrations. Thus, l.OxlO *7 M PMA was used in the rest of the experiments.

3.3.2 Kinetics of Static T-Flask Culture

Figure 3.5 shows the kinetics of PGE 2 production, glucose consumption, and

lactate accumulation in T-flasks. The PGE 2 concentration reached a maximum of 2.39

ng/ml at 40 hours. The production of lactate and PGE 2 appeared to be parallel to the

glucose consumption. At the end of culturing (56 hours), the medium pH decreased to

6.0. However, most cells still looked healthy under the microscope. The productivity of

FGE2 was 0.0548 ng ml *1 h*1, the glucose consumption rate was 0.11 g l *1 h' 1 and the productivity of lactate was 0.065 g I ' 1 h*1, respectively, in the first 8 hours.

3.3.3 Kinetics of Microcarrier Spinner Flask Culture

Figure 3.6 shows the concentration of PGE 2, glucose, and lactate in the microcarrier spinner flask culture. PGE 2 production continued even after glucose 49 consumption and lactate production had leveled off. It reached a maximum concentration of 8.67 ng/ml at 60 hours. The productivity of PGE 2 was 0.184 ng/mlt and the glucose consumption rate and the lactate accumulation rate was 0.28 and 0.121 g/L, respectively, in the first 10 hours. The final cell density in the microcarrier spinner flask was 9.26xl06 cells/ml, which was 4.1-fold of that in the static flask due to much more surface area supplied by microcarriers for cells to attach. The PGE 2 productivity was 2.35-fold and the PGE2 concentration was 2.64-fold of those in the T-flask. The pH in the spinner flask decreased to below 5.5 at 60 hours and at this time the dead cell number in the medium started increasing.

3.3.4 Selection of Bioreactor Packing Materials

Cells attachment conditions on various fibrous matrix were studied with scanning electron microscopy. As can be seen in the SEM photos (Figure 3.7), fibroblast cells had good attachment on 50% cotton/50% polyester fibrous matrix. However, the other types of the fibers did not show good attraction to the fibroblast cells. Hence, 50% cotton/50% polyester fiber was chosen as the packing material for use in the fibrous bed bioreactor.

3.3.5 Kinetics in a Batch Mode Fibrous Bed Bioreactor

When the fibroblast cells were cultured in the fibrous bed bioreactor under batch conditions, the PGE 2 production reached a maximum concentration of 18.98 ng/ml at

28.9 hours. At this time, the glucose concentration had decreased to 0.28g/l. The glutamine concentration in the medium had dropped to 0.238 mM at 8.7 hour and was completely gone by 20.7 hours. With the depletion of these substrates and accumulation of lactic acid and other toxic metabolites, the rate of prostaglandin E 2 production decreased quickly to almost zero at 28.9 hours. At the end of the batch culture, the pH in 50 the bioreactor was below 5.0. PGE 2 production could be improved if more glucose and

glutamine were present in the medium.

The productivity of prostaglandin E 2 (PGE2) was 1.26 ng ml ' 1 h*1, glucose

consumption rate was 0.38 g I *1 h* 1 and the lactate generation was 0.22 g l ' 1 h '1,

respectively, in the first 8.7 hours. PGE 2 productivity of the fibrous bed bioreactor was

about 23.0 folds and the PGE 2 concentration produced was 7 folds of those in the static

T-fiasks.

3.3.6 Kinetics in Well-Mixed Bioreactor

In a continuous well-mixed fibrous bed bioreactor, the continuous feeding relieved the stress of accumulation of toxic metabolites in batch culture. Figure 3.9 shows the kinetics of PGE2 production under continuous, well-mixed conditions. The concentration of PGE 2 reached a maximum of 22.32 ng/ml at retention time of 58.9 hours. At that retention time, the glucose concentration was 0.038 g/1, and glutamine in the medium was totally depleted at retention time 17.2 hours. If the glutamine level in the medium was increased, higher PGE2 concentration could be expected. The productivity of PGE 2 was 1.57 ng m l ' 1 h*1, glucose consumption rate was 0.35 g l ' 1 h*1, and the productivity of lactate was 0.26 g l ' 1 h '1, respectively, at 9.2 hours retention time.

3.3.7 Kinetics in Plug-Flow Bioreactor

The prostaglandin E 2 production in the fibrous bed bioreactor under plug-flow condition is shown in Figure 3.10. A high PGE 2 concentration of 23.48 ng/ml was attained at retention time 56.8 hours. With 10.5-hours retention time, the productivity of

FGE2 was 1.69 ng ml ' 1 h’1, glucose consumption rate was 0.36 g I ' 1 h*1 and the productivity of lactate was 0.20 g I *1 h 1, respectively. At the same retention time, the glutamine in the medium was totally depleted. Because of the depletion of the nutrients. 51 accumulation of the toxic metabolites, and decrease of pH in the bioreactor, the

productivity decreased significantly at 23 hours retention time and higher.

3.3.8 Comparison of Various Culture Systems

The comparison of PGE 2 production in various culture systems is shown in Figure

3.11. It is clear that the fibrous bed bioreactor had much higher PGE2 productivity than

the other two conventional culture systems. The higher productivity could be attributed

to the high cell density in the fibrous bed bioreactor, which was found to be 9xl0 7

cells/ml at the end of this study.

By comparison PGE 2 production by fibroblast cells at first 10 hours, the ratio of

productivity in different cell culture system was: plug flow bioreactor: well-mixed

bioreactor: batch bioreactor : microcaniers : static flask = 30.8 : 28.7 : 23.0: 3.4 : 1. The

ratio of highest PGE 2 concentration was: 9.84 : 9.36 : 7.97 : 3.6 : 1, respectively. Table

3.1 lists the PGE2 productivity, PGE 2 concentration, cell density, and specific PGE 2

productivity attained in various culture systems.

As can be seen in Table 3.1, the specific PGE 2 productivity in the fibrous bed was almost the same as that in the microcarrier spinner flask and was about 77% of that in the

T-flask. This indicated that the cell viability in the fibrous bed bioreactor was higher than

77% and most viable cells were healthy and functioned well. Thus, mass transfer limitation in the fibrous matrix was not significant even though there were high density of cells in the fibrous bed.

3.3.9 Kinetics of Substrates, Metabolite and Product

The concentration of lactate and PGE 2 are plotted against the glucose concentration. As shown in Figure 3.12, lactate yield from glucose was almost the same

(-0.575 g/g) in all culture systems. Figure 3.13 and Figure 3.14 show prostaglandin E 2 52 yields from glucose in various systems. The yield of PGE 2 from glucose in the static T-

flask was constant, - 0.72 Hg/g. for the microcarrier spinner flask culture, the PGE 2 yield

from glucose was also at -0.72 Hg/g, but it increased to -7 ^tg/g when the glucose

concentration was lower than 1 g/1. The dramatic change in PGE 2 yield indicated the cell

physiology change caused by the depletion of glutamine in the culture medium. The

FGE2 yield in the fibrous bed bioreactor under various operation modes were almost the

same, - 11.81 Hg/g, which was much higher than those in the T-flask and the

microcarrier spinner flask. The yield of prostaglandin E 2 from glutamine in the fibrous

bed bioreactor was almost constant, -6.5 (ig/mM before glutamine depletion (Figure

3.15).

3.3.10 Scanning Electron Micrographs

Figure 3.16 shows SEM pictures of NIH 3T3 fibroblast cells in the fibrous matrix.

From these pictures, two clearly distinct cell colony morphologies existed. In some areas

in the fibrous matrix, cells grew in poly-layer pattern and formed a "tissue-like" bio-film

on the surface of the fibers (Figure 3.16a,b). This indicated that the cells became

progressively more able to form multi-layers and interlaced in an highly crowed environment. This is also an indication of progressive loss of the contact inhibition characteristics of normal cells .29 In most areas of fibrous matrix, fibroblast cells were distributed evenly and attached to the fiber surface (Figure 3.16c). These cells were attached on surface in the 3-dimensional fibrous matrix and had normal morphology, which facilitated cell-cell interactions and helped cells maintain the in vivo physiological properties. 53 3.3.11 Conclusions

The fibrous-bed bioreactor gives high viable cell density for the anchorage-

dependent cell culture. The unique packing structure of the fibrous bed supplies a large

ratio of surface area to the bioreactor volume for cells to attach on. The gaps between the

spiral wound fibrous layers facilitate mass transfer and increase the bio-film renewability.

The three-dimensional structure of fibrous matrix provided a conducive environment for cell growth and helped cells to maintain their normal morphology and physiology. The fibrous bed supported a much higher number of cells to grow in the bioreactor and resulted in a much higher PGE2 productivity than those found in the conventional T-flask and microcarrier culture systems. The fibrous-bed bioreactor also gave a good long-term stability. In this study, the fibrous bed bioreactor was operated continuously for more than three months without suffering any problems. 54

Table 3. 1. Comparison of various bioreactor systems for fibroblast cell culture

Culture PGE 2 P G E 2 Cell Specific PGE 2 System Productivity Concentration Density Productivity (ng ml*1 h*1) (ng/ml) (cell/ml) (ngml'1 h*1 cell*1)

Static T-Flask 0.0548 2.39 2.25XI06 2.44xl0-8 Microcarrier 0.184 8.67 9.26x106 1.99x10*8 Fibrous Bed bioreactor 9.00xl07 Batch 1.26 18.98 1.40x10-8 * Well-Mixed 1.57 22.32 1.74x10*8* Plug-Flow 1.69 23.48 1.88x10-8

* based on the final total cell density of 9.0xl07 cells/ml at the end of the bioreactor study 55

Figure 3.1 The spirally wound structure of fibrous packing in the bioreactor 56

D.O./pH Meier

Sample Port pH/D.O Probe .

Air/5% Medium 002 Recirculation Flask

Product

Figure 3.2. Schematic diagram of well-mixed fibrous bed bioreactor for fibroblast cell culture 57

D.O./pH Meter

pH/DO Sample Probe Port

I I ill

Medium Product Tank Spiner Air/5% C 02 Flask

Figure 3.3 A schematic diagram of plug-flow fibrous bed bioreactor for fibroblast cell culture Figure 3.5. Kinetics of fibroblast cells N1H 3T3 in static flasks static in 3T3 N1H cells fibroblast of Kinetics 3.5. Figure PGE2 (ng/ml) 3 pGE2 (ng/ml) .4. Determination of optimal PM A concentration for cell stimulation cell for PM concentration A optimal of Determination .4. 0 2 3 1 4 5 0 2 1 3 4 0 0 020 10 10 20 l« (hr) Tlm« l« (hr) Tlm« 30 30 40 050 40 060 50 PQE2 60 70 0 3 u o o m 58 Figure 3.6 Kinetics of fibroblast cells NIH 3T3 on microcarriers on 3T3 NIH cells fibroblast of Kinetics 3.6 Figure PGE2 (ng/ml) 10 4 0 2 6 8 0 10 20 i* (hr) Tim* 050 30 40 Lactata PGE2 60 70 0 3 o e 10 O 59 60

t . , ' (i IID21 % 50% polyester/5 0% cotton

100% polyester 50% polyester/50% rayon

SI

100% glassflber ®*3 100% cotton

Figure 3.7 SEM pictures of fibroblast cells cultured with various fibrous materials 61

2 5

PGE2 c 4 < ► 20 o

IB •» c Lactate(g/I) u c U i o e> o o.

Gln(mM) 0 » 0 10 20 3 0 40 5 0 60 Time (hr)

Figure 3.8 Kinetics of fibroblast cells NIH 3T3 in the fibrous bed bioreactor at batch mode

5 2 5 PGE2

4 20 o

3 Lactate(g/I) 0) o CM c 2 UJ o o o & 1

(mM) 0 0 10 20 3040 5 0 6 0 Retention Time (hr)

Figure 3.9 Kinetics of fibroblast cells NIH 3T3 in the fibrous bed bioreactor at well-mixed mode

2 Figure 3.11 Comparison of PGE of Comparison 3.11 Figure PGE2 (ng/ml) 3 Concentration 3.10 Kinetics of fibroblast cells NIH 3T3 in the fibrous bed bioreactor bed fibrous the in 3T3 NIH cells fibroblast of Kinetics 3.10 0 2 1 3 4 5 0 1 2 3 4 5 6 7 8 90 80 70 60 50 40 30 20 10 0 n(mM 10 V eeto Tm (hr) Time Retention 20 in various culture systems culture various in ie (hr) Time at plug-flow mode plug-flow at F ib rous rous ib F 04 50 40 30 Glucose(g/l) ' ' 71 r 771 I , ^ 1'“ '— 1 2 3 PQE2 production by fibroblast cells cells fibroblast by production Batch * ELtd Well-mixed ♦ o Microcarriers T-flask Plug-flow Blorcacto

60 20 25

D. O Ul o> E 62 PGE2 (ng/ml) Lactate (g/l) 10 0 2 4 6 8 0 2 1 3 Figure 3.12 Lactate yield from glucose in various systems various in glucose from yield Lactate 3.12 Figure Figure 3.13 Prostaglandin E Prostaglandin 3.13 Figure 0 0 1 1 flask and microcarriers spinner flask spinner microcarriers and flask lcs (g/l) Glucose lcs (g/l) Glucose 2 2 • Plug-flow reactor Plug-flow reactor • Well-mixed • reactor Batch Microcarriers ■ flask • Static ■ T-flask • Microcarrier ■ 5 3 4 3 2 yield from glucose in static static in glucose from yield 4

5 63 64

Batch • Well-mixed Plug-flow

1 2 3 Glucose (g/l)

Figure 3.14 Prostaglandin E 2 yield from glucose in fibrous bed bioreactor

20

■ Plug-llow f ♦ Well-mixed ^ 15 ■ E Batch o* 10 04 UJ O a

x X X 0.0 0.5 1.0 1.5 2.0 Glutamlns (mM)

Figure 3.15 Prostaglandin E 2 yield from glutamine in the fibrous bed bioreactor Figure 3 16 Scanning electron micrographs of fibroblast cells NIH 3T3 in fibrous matrix 66

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PRIMARY LUTEAL CELL CULTURE

4.1 Introduction

The corpus luteum (CL) is a transient endocrine gland formed by differentiation

of follicular cells within the ovary after ovulation. If an embryo is not present, the corpus

luteum will regress and a new corpus luteum will form after the next ovulation. The

primary role of the corpus luteum is to synthesize and secrete progesterone, which is

essential for the maintenance of pregnancy (Smith, 1986). Development, maintenance,

and regression of the corpus luteum have been investigated for many years. However,

endocrine and cellular mechanisms regulating progesterone synthesis and secretion are

still unclear (Smith, 1986). The corpus luteum can be enucleated from large animals, and

after dissection, a large number of luteal cells can be attained. Although luteal cells can

synthesize cholesterol de novo, this cholesterol source is not sufficient to meet the

demands of steroidogenesis. Therefore, luteal cells are largely dependent upon

extracellular cholesterol sources, and large amounts of cholesterol are required to serve as

the substrate for maximal steroidogenic production (Pate and Condon, 1982).

Progesterone production is regulated and maintained by luteotropic and luteolytic

mechanisms, or by the pituitary hormone, luteinizing hormone (LH) (Hansel et al., 1978;

Rothchild, 1981). The specific mechanisms by which luteinizing hormone stimulates progesterone synthesis are currently unclear. A relatively reasonable model of second messenger stimulation was proposed by Niswender et al. (1980).

70 71 The physiological properties of luteal cells in vitro have been studied for a long

time. It was found by Pate and Condon (1982) that the presence of serum in the cell

culture medium inhibits the responsiveness of luteal cells to LH. Pate et al. (1987) also

found that the capability of production progesterone by luteal cells decreased with time

even in culture with the addition of LH. Many assumptions have been proposed to

explain this phenomenon. Besides the lower cell density compared with that in vivo (Pate

et al., 1987), Chow and Poo (1982) reported a redistribution of myotomal cell surface

lectin receptors when cell-cell contact occurred. Alterations in the arrangement of these

receptors could certainly change response to the hormone.

It is well documented that mammalian cells cultured in vitro will either retain

their morphological and biochemical differentiation and, thus, their protein synthetic and

secretory capabilities, or will lose them depending on the chemical composition and

structure of the culture supporting substrate (Singhvi et al., 1994). An impressive body of

literature has emerged in the last decade demonstrating that cell shape controls gene expression of extracellular matrix (ECM) proteins (Ben-Ze'ev, 1987), mRNA stability

(Mooney et al., 1991), post-translational modifications (Kabat et al., 1985), and differentiated functions of cells (Bissell, et al., 1987; Glowacki et al., 1983; and Lee et al.,

1984). Since there are evidences showing that substratum morphology alters the extent of cell spreading and cell shape (Hong and Brunette, 1987; Rovensky et al., 1971), it has been hypothesized that substratum morphology affects cell growth and function by modulating cell and nuclear spreading. Thus, it is important to consider two factors in developing an optimal method for culturing mammalian cells and producing therapeutic proteins: (1) The correct material composition and (2) the proper shape of the substrate

(Yang and Nandi, 1983). According to Yang's conclusions, both the shape of a cell and its orientation to neighboring cells are important in modulating the proliferative response of a cell to mitogen. Forkman and Moscona (1978) and Gospodarowicz et al. (1978) 72 have shown that for cells, especially primary cells, to express a fully differentiated

phenotype, a three-dimensional configuration is required. Hong and Brunette (1987)

cultured diploid epithelial cells on V-shaped parallel grooved titanium substrata. They

showed that secretion of proteinases by these cells on the grooved (3-D) substrata was

significantly higher than that on smooth flat substrata. To achieve a tissue-like three-

dimensional configuration, osteoblasts have been grown on a variety of matrices such as

alginate beads (Majmudar et al., 1991), Cytodex microcarriers (Sautier, 1992), and

collagen matrices (Casser-Bette et al., 1990). Such three-dimensional growth studies

have demonstrated cell differentiation to osteocytes, osteocalcin production, and site

specific mineralization, which were also compared with control osteoblasts grown on

two-dimensional surfaces, such as tissue culture plates (Hillslay and Frangos, 1994).

It has been highly desirable to employ a culture system in which the function of

cells is close to that in vivo. Gospodarowicz and Gospodarowicz (1972) reported that

bovine luteal cells were successfully cultured in vitro with static flasks. Stoklosowa and

Stadnicka (1973) successfully grew porcine and rat luteal cells, and Goldsmith and his

colleagues (1981) developed a continuous culture system for human luteal cells. Almost

all the experiments so far have been performed only in static flasks. However, these

systems had some problems and were not extensively used to study luteal steroidogenesis

in vitro. The disadvantages of culturing luteal cells in flasks are: (1) poor mass transfer

because the diffusion is the only mass transfer pattern in the static culture (Perry and

Wang, 1989), (2) nutrient depletion and accumulation of toxic metabolites at the end of the culture, which are deleterious to cells, and (3) alteration of cell morphology, because

the surface of static flasks can provide only a two-dimensional growth environment for cells to grow. According to Grinell (1983), flattening of the central mass of cells cause reorganizing of cytoskeletal elements within the wall and affects the cell's differentiated functions. It has also been clearly demonstrated by Emerman and Pitelka (1977) and 73 Ruzicka (1986) that cells cultured on flat plastic surfaces or on two-dimensional collagen sheets form a confluent epithelial sheet but more or less lose their differentiated ability for protein synthetic and secretory activities.

In this work, a fibrous-bed, immobilized cell bioreactor (Yang et al., 1992a,

1992b, 1994) was studied for cultivation of luteal cells. The fibrous bed bioreactor contained a spiral wound fibrous sheet for cell immobilization, with significant spaces between wound layers to allow uniform distribution and fast mass transfer of nutrients and products. The fibrous bed bioreactor provided not only a large surface area for cell attachment, but also three-dimensional environment which allowed cells to maintain their native morphology and physiology. The main objective of this study was to demonstrate that the fibrous bed bioreactor was a better in vitro culturing system than conventional static culture flasks in maintaining progesterone production by luteal cells. Luteal cells were cultured in the fibrous bed bioreactor which supplied a three-dimensional packing matrix for cells to attach and a fast mass transfer rate of nutrients and metabolites. The progesterone production by luteal cells in various systems were monitored. Cell migration and morphology changes in static culture and bioreactor were observed and analyzed.

4.2 Materials and Methods

4.2.1 Dissociation of Corpus Luteum (CL)

The dissociation of the corpus luteum (CL) followed the procedure previously used by Pate and Condon (1982). Corpora lutea were obtained from regularly cycling dairy cattle located at the Department of Dairy Science, The Ohio State University dairy farm. After enucleation the corpus luteum was put into Ham's F12 (Sigma) culture medium at 4°C. Then, the CL was cut into small pieces in a sterile petri dish and was further minced to small pieces in a small beaker. The minced tissue was placed in a 74 dissociation vessel with 20 ml Ham's F-12 and was stirred at 37°C for 10 minutes to wash tissue fragments. The washed tissue was allowed to settle down and the supernatant was discarded. Twenty ml Ham's F-12 containing 0.5% BSA and 0.25% collagenase was added to the remaining tissue to dissociate it for 45-60 minutes at 37 °C. The supernatant was collected and centrifuged. The pellet was resuspended in fresh medium (without collagenase). The aforementioned two steps were repeated for several times until most cells in the tissue were dissociated. The pooled cells were washed 2-3 times with Ham's

F-12 (without collagenase) and resuspended in 5-10 ml medium and the viable cells were counted using trypan blue on a hemocytometer.

4.2.2 Culture Medium

Cells were cultured in Ham's F-12-Hepes culture medium (containing 4.5 g/l glucose, K. C. Biological, KS) supplemented with 5 Hg/ral insulin, 5 (ig/ml transferrin, and 5 ng/ml selenium (ITS premix, Collaborative Research, Lexington, MA). Penicillin

(100 U/ml)-streptomycin (100 |ig/ml) (K.C. Biological) and gentamicin (20 ng/ml. Grand

Island Biological Co., Grand Island, NY) were also added. The serum free medium used in this experiment did not allow any fibroblast cells, if present, to survive in the culture systems.

4.2.3 Cultivation in Static T-Flasks

T-125 culture flasks were used in these experiments. Before the cells were transferred, medium with 10% serum was used to coat the inner surface of the flasks for

2-3 hours, which helped cells to attach to the surface. The fresh medium was then used to wash the remaining serum in the flasks away. Culture medium with ITS (insulin, transferrin, and selenium) was placed into the flasks. About 6x10* cells were transferred 75 into each flask containing 200 ml of the medium. The flasks were incubated in a 5% CO 2 incubator (Napco E series, Model 302) at 37°C.

The effect of LH stimulation on progesterone production by luteal cells were studied first. Two T-flasks were used in the experiment. The one without LH addition was designated as a control and the other with LH addition was used to test cell response in progesterone production. Twenty four hours after the cells were introduced to the T- flasks, the medium in the flasks was changed with fresh medium and every 48 hours thereafter. With each medium change, LH (100 ng/ml) was added, and samples were taken from the flask one hour after LH addition. The samples were frozen for future glucose and progesterone analyses.

The effect of medium change frequency on progesterone production was also studied with three T-125 flasks. The medium in the culture flasks was changed at 12,24, and 48 hour intervals, respectively. Every 48 hours, 100 ng/ml of LH was introduced into the individual flask. One hour after LH stimulation, samples were taken from each flask and frozen for glucose, lactate, and progesterone analyses. All experiments were either duplicated or repeated once and the average from the two sets of data were reported.

4.2.4 Cultivation in Fibrous Bed Bioreactor

4.2.4.1 Bioreactor Construction

The immobilized cell bioreactor was made of a glass column with 4.5 cm diameter and 15 cm height. Water of constant temperature was circulated through the water jacket around the bioreactor to keep the temperature in the reactor constant. A piece of 50% polyester/50% cotton Tiber (100 x 10 x 0.1 cm) was spirally wound and put into the bioreactor on top of 1" thick 1/4" glass beads, and the glass beads were used to help form an even distribution of the input medium. The height of the fibrous packing 76 was 10 cm and the working (liquid) volume of the reactor was about 200 ml. Two

peristaltic pumps (Masterflex, Cat.#N-07520-14) were used in the system. One pumped

the fresh medium into the bioreactor, and the second one recirculated the medium in the bioreactor through the recirculation flask at a high flow rate (80 ml/min) to provide a well-mixed condition in the reactor. A pH probe (Cole Parmer, #G-5662-10), a dissolved oxygen probe (Cole Parmer, #G-5644-00), an inoculation port, and a sampling port were installed on the top of the reactor to monitor the pH and dissolved oxygen in the reactor.

A spinner flask (with less than 100 ml liquid volume) was installed in the recirculating loop. Sterilized air containing 5% carbon dioxide were used to flush the surface of the liquid in the recirculation flask to supply the oxygen to the medium and balance the pH in the bioreactor. The effluent overflow out of the spinner flask by the pressure built up inside the system. Figure 4 .1 shows the schematic diagrams of the fibrous bed bioreactor with spirally wound packing structure used in this work.

4.2.4.2 Selection of Packing Materials

Five different fibrous materials were used to test the attachment condition of luteal cells and to select the appropriate packing materials for use in the fibrous bed bioreactor. The fibrous materials tested were: (1) 100% polyester, (2) 100% cotton, (3).

50% polyester/50% cotton, (4) 100% fiber glass, and (5) 50% polyester/50% rayon. All different fibers were cut to small pieces (lxl cm) and placed in a 24-well culture plate.

1.0 xlO 6 cells were transferred into each well and the culture medium was changed every

24 hours. After 2 days culture, the fibrous materials were taken out and processed for scanning electron microscopy (SEM). According to cell attachment condition, an optimal fiber material was chosen for the rest of the experiments. 77 4.2.4.3 Cell Immobilization

The bioreactor system was autoclaved at 121 °C for 1 hour at least twice at a 24

hour interval. After the bioreactor cooled down to the room temperature, the water bath

connected to the water jacket of the reactor was started to keep the temperature of the

bioreactor at 37°C. The medium with 10% fetal bovine serum was pumped into the

reactor until the reactor was full of the medium. The packing fibers and the inner surface

were coated with serum for 12 hours to aid cell attachment. Then, 1 liter of filter-

sterilized medium was prepared in a 2-liter flask and pumped into the bioreactor to wash the serum away. Then, culture medium with ITS (insulin, transferrin, and selenium) was pumped into the bioreactor. About 4.8xl0 9 cells were injected from the inoculation port by using a 10-ml syringe. The recirculation was turned on at a slow speed (20 ml/min) to ensure that cells were evenly distributed in the fibrous packing. After 4-5 hours the medium became clear, indicating that most cells had been attached to the fiber surface.

Three bioreactors were studied under batch, fed-batch, or continuous culture condition, respectively. These experiments were also repeated and the average result from the duplicates were reported.

4.2.4.4 Cultivation under Batch Mode

The 200 ml media in the bioreactor were replaced with fresh medium every 48 hours. Right after the medium change, a small volume of concentrated luteinizing hormone was introduced into the bioreactor to a final LH concentration of 100 ng/ml in the bioreactor. One hour after the stimulation, samples were taken from the bioreactor to measure the concentrations of glucose and progesterone. 78 42.4.5 Cultivation under Fed-Batch Mode

Every 12 hours 50 ml media in the bioreactor were replaced with fresh medium.

Every 48 hours, LH was added into the bioreactor to a final LH concentration of 100 ng/ml. One hour after the LH addition, samples were taken from the bioreactor. The samples were frozen for glucose and progesterone analyses.

4.2.4.6 Cultivation under Continuous Mode

The medium in the bioreactor was replaced with fresh medium after the first 24 hours of culture, and LH was injected into the bioreactor to a final concentration of 100 ng/ml. The culture medium was then pumped continuously into the bioreactor at a feed rate of 100 ml/day (-48 hours retention time). A final concentration of 100 ng/ml LH was added into the bioreactor every 48 hours. Samples were taken from the bioreactor 1 hour after LH was added. The samples were frozen for glucose and progesterone analyses.

4.2.5 Analytical Methods

4.2.5.1 Cell Density in the T-Flasks

The cells were removed from the surface of the flasks to which they attached with

0.125% trypsin. The resulting cell suspension was diluted with 0.1% Dulbecco's phosphate buffered saline (PBS) and counted on a hemocytometer.

4.2.5.2 Glucose and Lactate

The concentration of glucose in the samples was measured with either YSI glucose analyzer ( Model 2700 SELECT) or high performance liquid chromatography

(HPLC). The HPLC system consisted of a high pressure pump (Waters, Model 6000A), and injector (Waters, U 6 K), an organic analysis column (Bio-Rad, Model HPX-87H; 79 condition: 45°C), a column heater (Bio-Rad), a RI detector (Waters, Series 410

differential refractometer; condition: scale factor 2 0 , sensitivity 16, and internal oven

temperature of 4S°C), an integrator (Spectra Physics, Model 4270; conditions: chart

speed 0.25 cm/min, attenuation 32; minimum area 5000, peak height 1, and run time 20

minutes). 0.01 N H 2SO 4 was used as the eluent at 0.6 ml/min flow rate. Lactate concentration was determined by HPLC method described before.

4.2.5.3 Progesterone

The concentration of Progesterone (P 4) was analyzed with duplicates by enzyme- linked immuno-sorbent assay (ELISA) and radioimmunoassay (RIA) (Pate and Condon

1982). All measurements were duplicated.

4.2.5.4 Scanning Electron Microscopy (SEM)

After all experiments were finished, fiber samples (lxl cm) were taken from different positions of the packing. After leaving the samples in 2.5% formaldehyde solution overnight, they were rinsed 10 times with double distilled water, each time for

15 minutes. Then dehydrated with 20 to 70% ethanol in 10% increments by leaving them in different concentration solutions for 30 minutes each time. The next day, they were dehydrated with 80% ethanol and twice with 95% and 100% ethanol for 30 minutes each time. The samples were kept in 100% ethanol overnight again and then dried at critical point with liquid CO 2 in a Pelco Critical Point Drier in the Anatomy Department of the

Ohio State University. All steps, except that for critical drying, were carried out at 4°C.

SEM photos were taken by using JOEL-Model 820 SEM at Geology Department of the

Ohio State University. 80 4 J Results

4.3.1 Progesterone Production by LH Stimulation

Luteal cells in the fibrous bed bioreactor responded to LH stimulation with an

increase in progesterone. Figure 4.2 shows a typical profile of this response. One hour

after every stimulation, progesterone concentration rose drastically and then decreased

gradually to a low level close to zero. It is interesting that the consumption of glucose is

well correlated with progesterone production, suggesting that progesterone production

required glucose and other nutrients. At the end of every 48 hours, lactate concentration was approximately 0.7 g/L (not shown) and the pH of the medium was down to 6 .8.

4.3.2 Cultivation in Static T-Flasks

Figure 4.3 shows progesterone production with and without LH stimulation in static T-flasks at various culture times. Without LH stimulation, the progesterone production at day one was -254.1 ng/500,000 cells. The progesterone production then decreased significantly during the next 6 days of cultivation. The average rate of decrease was 20.76 (ng/50,000 cells/day). Then, it kept a relatively low level for almost two weeks and finally decreased to zero at day 19. With LH stimulation, the progesterone production was much higher, about 423.0 ng/50,000 cells at day one.

However, the progesterone production by luteal cells also decreased with cultivation time at a rate (24.86 ng/50,000 cells/day) similar to that without LH stimulation. The progesterone response decreased to 77.3 ng/50,000 cells at day 19.

The decline in progesterone production by luteal cells with cultivation time is an indication of decay in cell viability or function, which may be assumed to follow the first order reaction kinetics: 81

In S r - -k dt (2) Co

where C is the progesterone production (ng/50,000 cells), Co is the initial progesterone

production (ng/50,000 cells), kd is the progesterone decay rate constant, and t is the

cultivation time (day).

The effect of nutrient change frequency on the declining rate in progesterone

production by luteal cells in T-flask is shown in Figure 4.4. The decay rate constant, kd,

decreased with increasing the medium change frequency. The kd values for medium

changes at intervals of 12, 24, and 48 hours were 0.058 0.068 and 0.18 (1/day),

respectively. It is obvious that the decline of progesterone synthetic capability at 48-hour

medium change is much faster than those at both 12-hour and 24-hours. The glucose

level in the medium in 48-hr intervals was low (0.32 g/1) before the medium change.

However, the glucose concentrations in 12-hr and 24-hr tests were 2.74 and 1.87 g/1

respectively. These results show that nutrient supply might have played an important role

in progesterone secretion.

4.3.3 Cultivation in Fibrous Bed Bioreactor

4.3.3.1 Selection of Fibrous Materials

Attachment of luteal cells on various fibrous materials are illustrated in the

scanning electron micrographs (Figure 4.5). Apparently, luteal cells attached best to 50%

polyester/50% cotton material. Cell attachment to 100% polyester and 50% polyester/50% rayon materials was not as good, but was somewhat better than 100% cotton. Therefore, the fiber material with 50% polyester/50% cotton was chosen as the packing material in the bioreactor to conduct the following experiments. 82 It was also noticed that cells grew well on the surface of a cover glass. However,

since they attached to a flat surface, most cells could only maintain 2 -dimensional

morphologies, which were different from those of cells in the fibrous matrix.

4.3.3.2 Cultivation of Luteal Cells in the Fibrous Bed Bioreactor

The progesterone production by luteal cells in the fibrous bed bioreactor under

various operation conditions was shown in Figure 4.6. In general, progesterone

production by luteal cells in the fibrous bed bioreactor also decreased with cultivation

time, but the declining rate was much slower than that in the T-flask (Figure 4.7). The initial progesterone production was about the same for both T-flask and fibrous bed bioreactor with LH addition (Figure 4.7a)

Figure 4.7b shows that the decay rate was much smaller for cultures in the fibrous bed bioreactor than that for the T-flask cultures. Table 4.1 shows the comparison of initial and final progesterone production by luteal cells in response to LH stimulation and the declining rate during the first 5 days and the k

4.3.3.3 Cell Morphologies and Attachment Conditions

The morphologies of luteal cells in T-flask at different cultivation times are shown in Figure 4.8. After the cells were immobilized on the surface of the static flask, the cells formed evenly distributed cell pavement on the surface of the static flasks. By day 7, clumps of cells appeared to interrupt the cell pavement. Between 2 to 3 weeks in culture, the luteal cell population showed signs of stress; holes appeared on the pavement, and cell sloughing was observed. The cells were isolated or in groups to form flat, two- dimensional, discontinuous sheets.

Cells seeded in the bioreactor adhered to the fibers quickly and by 4-5 hours formed an almost confluent luteal cell pavement. There were cells beneath the surface 83 layer, and some invaded into the fiber bundles. The diameter of cells in the bioreactor

was about one third of that in the T-flasks. The morphologies of luteal cells in the fibrous

bed bioreactor at different cultivation times are shown in Figures 4.9. After 4-5 days 1 culture, cell migration was observed. The evenly distributed cells moved slowly toward certain areas where cell aggregation occurred. Between day 15 to day 20, large cell clumps were formed and the aggregated cells were dispersed in different areas on and within the fibrous matrix.

4.4 Discussions

There has been little success in maintaining the differentiated functions of primary cells when they were cultured in static T-flasks. Pate and Condon (1982) observed that the capability of progesterone production by luteal cells decreased with the culturing time even with the addition of LH. Primary cell cultures of lactating bovine mammary epithelium show a decline in specific protein synthesis (Forsyth and Jones, 1976), even in the presence of lactogenic hormones (Anderson and Larson, 1970). There are several possibilities to cause the lack of secretory differentiation by luteal cells in static, two- dimensional culture.

4.4.1 Cell Interactions

When the luteal cells were cultured in the static flasks, the cells spread out on the inner surface of the flasks and became relatively flat. Although they were joined by tight and adhering junctions, this situation, in which the lateral cell surface were not in as close contact with cell neighbors as was found in vivo, may have result in loss of cell interactions necessary for secretory expression. Pate et al. (1987) reported that LH- stimulation of progesterone production by luteal cells was dependent upon the interaction of LH molecules with a specific membrane receptor. It was also found by Jones et al. 84 (1992) that the membrane receptors lost during in vitro culture, which may partially cause

the decrease of the response of luteal cells to LH.

The formation of the aggregated luteal cell clumps in both static flasks and fibrous

bed bioreactor was an interesting phenomenon. The migratory behavior of a number of

cell types on three-dimensional collagen gels has been described in a number of papers

(Schor, 1980; Schor et al. 1982). According to Miyoshi et al. (1994), cell aggregation is

generally considered to improve and preserve the metabolic functions of individual cells.

This change of cell configuration was expected to favor metabolic functions. Many

investigators reported that in primary cells, such morphological changes occurred 2-5

days after inoculations (Koide et al. 1989; Landry et al., 1985; Takahashi et al., 1989). In

fibrous bed bioreactor, the packing matrix seemed to facilitate the cell migration and

aggregation by providing a three dimensional environment, which could not be

accomplished in two dimensional static flasks.

4.4.2 Cell Morphology

When luteal cells were dissociated from the corpus luteum and cultured in static flasks, they experienced dramatic changes of living environment. Firstly, a three- dimensional geometry growth condition was changed to two-dimensional, which caused the cell morphology to change; secondly, specific cell interaction characteristics of the histology of the tissue were lost, and often the biochemical properties associated with it.

The morphologic change of the cell resulted in the reorientation of membrane geometry, cell structure, and organelles in cells, which might cause changes in cellular DNA synthesis and the process of mitosis, and other biochemical and physiological properties.

The loss of specific cell interaction characteristics of the histology of the tissue might have the same effect 85 The fibrous bed bioreactor provided a suitable environment for significant specific differentiation of luteal cells. This environment satisfied the requirement of luteal cells to attach to a substrate while at the same time permitting them to aggregate without severing them attachment, since the fibers were not as rigid as static flasks. The plasticity of the fibers permits the luteal cells to change shape; the contraction of the cell diameter reflected this change. The cells resemble their native shape and achieved a configuration much more tissue-like than in a culture spread out on plane surface and perhaps more conducive to expression of the differentiated state.

4.4.3 Nutrient Factors

According to Emerman and Pitelka (1977), another consideration is the accessibility of nutrients to the cells. The basolateral surface of the cells received nutrients and hormones in vivo; necessary receptors and transport mechanisms may be concentrated on these surfaces. When cells adhered to plastic, impermeable substrates, presumably some nutrients and growth factor are prevented by occluding junctions from entering spaces below the apical surface of the luteal cell pavement.

After LH binds to the membrane receptor in the plasma membrane the steps involved in signal transfer, enzyme activation, and pregnenolone and progesterone synthesis require energy. Although many nutrients are added to the culture, they are still not rich enough compared with that in vivo, especially when luteal cells are cultured in serum free medium, which makes cells to adjust certain metabolisms to survive themselves in the new environment. In addition, when cells are cultured in vitro, either in static flasks or fiber matrix, unlike that in vivo, they have to expend energy to attach to the surface of the flasks or fibers to form a monolayer, which also demands energy. 86 4.4.4 Mass Transfer Rate

It was found that the consumption rate of glucose and other nutrients were well correlated to progesterone production and the cells in the bioreactor had a higher response to LH than those in the static flasks at the same culture time. When cells are cultured in static flasks, the only pattern of mass transfer is natural diffusion and with culture time passing, the nutrients close to cells are depleted and toxic metabolites accumulate around the cells. If the diffusion is not faster than the reaction, the process becomes mass- transfer limited and the cells are in an unfriendly micro-environment.

The fibrous bed bioreactor allowed exposure of the cells to everything presented in the medium. To improve the mass transfer rate, a perfusion system was used that was more efficient than the static process. Therefore, in the perfusion bioreactor, because the medium is recycled and replenished and the mass transfer condition is improved, the nutrients were supplied to and the toxic wastes were removed from the cells at a higher efficiency, which relieved the environmental stresses on the cells and helped cells to maintain their differentiated functions. 87

Table 4.1. Comparison of progesterone production by luteal cells in response to LH addition in various culturing systems

Culture Initial P 4 P4 Declining Final P 4 kd System Response Rate Response (ng/50,000 cells) (ng/50,000 celk/d) (ng/50,000 cells) (d'1)

Static T-flask without LH 254.1 20.8 ~ 0 0.090 with LH 423.0 24.9 77.3 0.086 Fibrous Bed Bioreactor Batch 394.90 6.15 394.9 0.036 Fed Batch 522.0 9.2 325.0 0.033 Continuous 513.8 5.2 411.8 0.015 88

D.O./pH Meter

Sample Port pH/D.O Probe r

,

Air/5% 002 Recirculation Flask

Product

Figure 4.1. Schematic diagram of well-mixed fibrous bed bioreactor for luteal cell culture Figure 4.2 Progesterone production by luteal cells in response to LH addition LH to response in cells luteal by production Progesterone 4.2 Figure Progesterone Production (ng/50,000 cells) 1000 400 0 - 600 BOO ‘♦-J «— J - ♦ ‘ 0 60 - 20 -2 0 -4 0 -6 in the fibrous bed bioreactor bed fibrous the in lue ie (hour) Time ulture C d LH Add 0 20 40 d LH Add Glucose 60 rgseoe _ Progesterone 80 100

Glucose (g/L) 89 600

without LH o =3 4 0 0 with LH

8 3 200

6 7 8 9 10 11 12 13 14 15 16 17 18 1920 Culture Time (day)

Figure 4.3 Progesterone production by luteal cell cultured in Static T-flasks

-1

o O U

-2 12 hr. Interval 24 hr. Interval 48 hr. Interval

0 4 8 12 16 20 24 28 Time (day)

Figure 4.4 Effect of medium change frequency on progesterone production 91

50% polyester/50% cotton

50% polyester/50% rayon 100% polyester

100% glassflber 100% cotton

Figure 4.5 SEM photos of luteal cells attachment to various fibrous materials 800 t— r -i— i— r — i— i— i— r batch a 700 o fed batch •N continuous O 600 T33 —£ £O u o 500 o* o a °. 400 P ° ii v . 300 a)» a 2P oo w 200 - uo 0. 100

6 7 8 9 10 11 12 1’ 14 15 16 17 18 Culture Time (day)

Figure 4.6 Progesterone production by luteal cells cultured in fibrous bed bioreactor o o .1 1 3 7 1 1 1 1 1 21 19 17 15 13 11 8 7 5 3 1 Figure 4.7 Comparison of progesterone production by luteal cells to LH to cells luteal by production progesterone of Comparison 4.7 Figure Progesterone Production (ng/50,000 cells) 600 400 500 200 100 0 ia (day) Tima 2 4 in various systems various in 6 lue ie (day) Time ulture C 8 (a) b) (b iratr (fed-batch) Bioreactor • 0 2 4 6 8 20 18 16 14 12 10 oeco (conlinuoua) ioreactor B ai Flaak tatic S Sai Pak wLi • (w/LJi) Plaak Static ■ a a oeco (batch) ioreactor B ♦ T-Flask (no LH) (no T-Flask ♦ b a Reactor (Batch) Reactor a (Continuous) eactor R ■ eco FdBatch) (Fed-B Reactor T -Flask (with LH) (with -Flask T ° □ - w/ LHP\ P H L /o (w 93 Figure 4.8 SEM photos of luteal cell morphologies and attachment condition in static T-flask 95

Figure 4.9 SEM photos of luteal cell morphologies and attachment condition in fibrous bed bioreactor 96

Figure 4.9 (continued)

ar---

> '■ i

DAY 20 97

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HYBR1DOMA CELL CULTURE

5.1 Introduction

One of the most important economic and social impacts of biotechnology is the

production of monoclonal antibodies by using mammalian hybridoma cells. Kohler and

Milstein (1975) demonstrated that individual clones of normal antibody secreting cells

could be immortalized by fusion with myeloma cells. Fusion resulted in formation of

hybridoma cell lines, which secreted the antibodies of the antibody forming cells and

could be propagated indefinitely in vitro like the myeloma cells.

The techniques of monoclonal antibody production by culturing hybridoma cells

have been studied for several decades and made great progress. Suspension cell systems

are frequently run on a batch basis, in which cells produce antibody for only a few days

before they die. This causes the difficulties in downstream process for removal of dead

cells from the antibody-containing medium. Suspension systems usually achieve relative

low concentrations of cell, 106 cells/ml in batch systems and 107 cells/ml in continuous systems (Altshuler et al., 1987; Katinger, 1987). Dalili and Ollis (1990) cultured hybridoma cell line MRC OX-19 in 100 ml T-flasks for 170 hours and the highest cell density and monoclonal antibody concentration was about 2X106 cell/ml and 21 mg/1, respectively. However, since batch production is a well-understood technique and easy to scale up, most large industrial productions are still using batch. Backer and his colleagues (1988) in Eli Lilly & Company reported that by using a marine

102 103 impeller agitator, they successfully cultured hybridoma cells in a 1300-liter fermenter.

Lambert et al., (1987) grew 35 different monoclonal antibody lines in 1000 liter airlift

system and achieved antibody concentration levels ranging from 40 to 500 mg/1.

Many modifications have been made to increase the cell density and productivity

of suspension bioreactors. Broise et al. (1992) coupled stirred tank to an external

tangential flow filtration, the cell density and MAb concentration increased to 8xl0 6

cells/ml and 120 mg/1, respectively. By using an air-lift bioreactor connected to a cell

settler for cell retention, Hulscher and his colleagues (1992) increased the MAb

productivity by a factor of 17 and the cell density by 4 compared with conventional batch

systems. Shi et al. (1992) designed a new impeller to improve the oxygen transfer in the

bioreactor, the total hybridoma cell concentration was increased to 3.4 xlO 7 cells/ml, and

the MAb product concentration was also increased to 512 mg/1. Increase of cell density

and MAb production by using a dialyzed continuous suspension culture system was also

reported by Linardos et al. (1992).

Most hybridoma cells can be cultured in conventional fermenters. However, it is

more ideal to culture the cells under immobilized conditions, because the hybridomas

require a stable, shear stress free physical and chemical environment to achieve optimal

growth and productivity (Emery et al., 1987). Immobilization techniques which seek to

make use of a more or less conventional reactor design involve in entrapping the cell

within some forms of particles, such as agarose (Nilsson, et al., 1983), alginate (Siacore,

1984), or a membrane-bound capsule (Rupp, 1985). The immobilized cells are maintained in a protected, biochemically conditioned environment. Beads and porous particles allow a cell-free product stream to be withdrawn from the reactor. In all of these cases, the reactors can easily be perfused with a continuous supply of fresh culture medium, extending the productive lifetime of the cells and increasing the cell concentrations obtained in the reactors. 104 Novel bioreactor designs which do not use particle-entrapped cells achieve cell

immobilization by enclosing the cells within one or more specific compartments.

Examples include many variants of hollow-fiber bioreactors (Altshuler, et al., 1987;

Tharakan et al., 1986), in which cells are enclosed within the shell of a hollow-fiber cartridge, and the membrane bioreactor design of Klement et al.(1987) and Scheirer

(1988), in which cell are contained in a thin layer between two planar membranes. These systems, too, maintain the cells in a conditioned and protected environment. Lee (1991) compared the monoclonal antibody production between suspension cell and immobilized cell culture by immobilizing S3H5/y2bA5 hybridoma cells in alginate beads. The experimental results showed that the viable cell density and the volumetric MAb productivity were three times of those in the suspension cell culture. Cadic and his colleagues (1992) immobilized hybridoma cells in agarose beads and achieved 20-fold more cell density than that in suspension culture. Broise et al. (1992) cultured the hybridoma cells in both alginate membrane and hollow fiber bioreactors and achieved 2 to 1.5 folds increase of MAb production, respectively. A better result was obtained by

Altshuler (1987) and his colleagues (4.5 and 7 folds, respectively) using a hollow fiber bioreactor. Hagenom and Kargi (1990) developed a coiled tube bioreactor and achieved a 6 folds increase of MAb increase compared with spinner flasks. New Brunswick

Scientific Co., Inc. (Wang et al., 1992; Biomass News, 1993) developed a packed-bed bioreactor (CelliGen) whose productivity could reach as 12-fold high as that in static and stirred suspension culture systems.

Inevitably, such improvements in bioreactor performance also raise some problems. The hollow fiber, membrane, and encapsulation bioreactors allow high cell densities to be maintained (about 10s cells/ml), but the scale up of these methodologies are usually problematic (Wang, et al., 1992). Cadic and Dupuy (1992) reported that the production of MAb decreased after two weeks cultivation due to the spherical colony 105 formation causing mass transfer limitations. Broise et al. (1992) also observed the

production of MAb started reducing after three-week cultivation because of the

destruction of the beads. The fiber-bed bioreactors suffer clogging and channeling

problems after a period of time of operation. Experiments by Wang et al. (1992) showed

that the production of MAb in the packed-bed bioreactor (CelliGen) started decreasing after 35 days cultivation. The reasons behind these phenomena are complicated.

However, the poor mass transfer plays an important role, especially after a large amount of cell mass building up inside the bioreactor.

In this work, the production of monoclonal antibody (MAb) IgG2b by mouse hybridoma cells HD-24 immobilized in the fibrous bed bioreactor was studied and compared with cultures in static T-flasks and spinner flasks. The MAb production, glucose and glutamine consumption, lactate production, and cell growth were evaluated.

The production level of MAb and cell density achieved in various culture systems were used to assess various bioreactor systems studied. The long-term stability of the fibrous bed bioreactor was tested for a period of over 30 days. Cell morphology and cell immobilization in the fibrous matrix was also studied using scanning electron microscopy and confocal microscopy.

S2 Materials and Methods

5.2.1 Cell Line

Mouse hybridoma cell line HD-24 was supplied by Dr. Richard Mortensen's Lab,

Department of Microbiology, The Ohio State University. The culture was maintained in

T-25 culture flasks in a CO 2 incubator (Napco E series, Model 302) at 37°C. 106 5.2.2 Medium

The medium consisted of 90% (v/v) DMEM (containing 4.5g/l glucose and 4mM

glutamine; Irvine Scientific, #9418-10718), 10% (v/v) fetal bovine serum (Cell Culture Laboratories #10-1010-50), and 60 |ig/ml gentamicin (Whittaker #1301-898-7025). The

medium was sterilized by filtration through a 0.22-^m medium filter (Coming #25970-

33).

5.2.3 Batch Culture Studies

5.2.3.1 Effect of Cell Seeding Density

Six different seeding cell densities (6.4x10s, 3.2x10s, 1.6x10s, 8x10*, 4X104, and

2X104 cells/ml) were tested. Cells were inoculated into six T-125 flasks at different

seeding densities and the flasks were incubated in the CO 2 incubator at 37 °C. Samples (1

ml) were taken every 12 hours from the T-flasks and viable cell densities were counted to determine the optimal cell seeding density.

5.2.3.2 Static T-Flasks

A total number of 8.0X106 cells were seeded into a T-125 static flask with 25 ml medium. Then, the T-flask was incubated in the incubator at 37°C. Samples (1 ml) were taken every 12 hours from the T-flask and were frozen immediately for future analysis of

MAb, glucose, lactate, and glutamine after the viable cell number had been counted.

5.2.3.3 Spinner Flask

Cells were seeded into a 500 ml spinner flask (Bellco) with 200 ml liquid medium to a final cell concentration of 3.2x10s cells/ml. The spinner flask was placed on a stirring plate (Bellco, Multi-Stir 4 low profile stirrer) with a stirring speed of 70 rpm and incubated in the CO 2 incubator at 37 °C. Samples (1 ml) were taken every 12 hours from 107 the spinner flasks and were frozen immediately for future analysis after the viable cell

number had been counted.

5.2.4 Fibrous Bed Bioreactor

5.2.4.1 Selection of Fibrous Materials for Bioreactor

Five different woven fibrous sheet materials were used to test the attachment

condition of hybridoma cells. They were: 100% polyester, 100% cotton, 50%

polyester/50% cotton, 100% fiber glass, and 50% polyester/50% rayon. The fiber

materials were cut to small pieces (lxl cm), autoclaved, and placed in a 24-well culture

plate. Cells with final concentration of 3x10 5 cells/ml were inoculated into each well and

the medium was changed every 72 hours. After seven days culture, the fibrous materials

were dehydrated for examination using a scanning electron microscope (SEM) to

evaluate attachment condition and to choose the fibrous material for the fibrous bed

bioreactor.

5.2.4.2 Bioreactor Construction

The fibrous bed bioreactor was made of a glass column with 4.5 cm inner

diameter and 19 cm length. Water of constant temperature (37°C) was circulated through

the water jacket around the bioreactor to keep the temperature in the reactor constant. A

piece of the fibrous sheet (1 0 x 100 x 0.1 cm) was spirally wound and packed in the

bioreactor on top of 1" thick 1/4" glass beads, which were used to help distribute the feed

medium. The height of the fibrous bed was 10 cm and the bioreactor working (liquid)

volume was about 200 ml. Two peristaltic pumps (Masterflex, No. N-07520-14) were

used in the system. One pumped the fresh medium into the bioreactor, and the second one circulated the medium in the bioreactor through the recirculation flask at a high flow rate (80 ml/min) to provide a well-mixed condition in the reactor. A pH probe (Cole 108 Parmer, #G-5662-10), a dissolved oxygen probe (Cole Parmer, #G-5644-00), an

inoculation port, and a sampling port were installed on the top of the reactor. A 2-liter spinner flask (containing 0.1 liter liquid) was installed in the recirculation loop.

Sterilized air containing 5% carbon dioxide was used to flush the surface of the liquid in the recirculation flask to supply the oxygen to the medium and to balance the pH in the bioreactor. The effluent overflow out of the spinner flask by the pressure built up inside the system. Figure 5.1 and Figure 5.2 show the schematic diagrams of the spirally wound structure of fibrous matrix and the well-mixed type reactor system used in this work, respectively.

The well-mixed bioreactor could be converted to a single-pass bioreactor with some modifications: ( 1). disconnecting the recirculation loop and converting the spinner flask to a feeding tank, in which the fresh medium was surface-aerated (95% air and 5%

CO2) before being pumped into the bioreactor, and (2). moving the effluent port to the top of the bioreactor. A schematic diagram of the plug-flow type bioreactor system is shown in Figure 5.3. In this single pass bioreactor, concentration and pH gradients existed along the reactor length. Since the feeding stream passing through the glass bead layer had the same concentration at all radial position and the flow velocity along the bioreactor length was small (1 cm/hr), the flow condition in the bioreactor was close to plug flow.

5.2.5 Bioreactor Start-up

The entire bioreactor system, including the pH and DO probes, the feed tanks, the recirculation flasks and all the tubing and connections, was autoclaved for 1 hour at

121°C, 15 psig at least twice, at a 24-hour interval. Then, 1 liter filter-sterilized medium was prepared in a 2-liter flask. After the bioreactor cooled down to the room temperature, the medium was pumped into the reactor until the reactor was full of the 109 medium. The water bath connected to the water jacket of the reactor was started running

to keep the temperature of the bioreactor at 37°C. Concentrated cell suspension from

centrifuging from T-flask culture was injected from the inoculation port into the

bioreactor by using a 10-ml syringe to a total cell density of 2 x 10 5 cells/ml. The

recirculation was turned on at a slow speed (2 0 ml/min) to ensure that the cells were

evenly distributed on the fibrous packing. No fresh medium was added into the

bioreactor until the pH dropped to 6 .8 or oxygen to 15%. Then, a low pump speed (0.2

ml/min) was set to replenish the medium in the bioreactor. The pump speed was

increased gradually until it reached 0.8 ml/min. It took a total of 20 to 30 days for the

bioreactor to reach pseudo-steady-state as determined from stable outlet glucose and

lactate concentrations. At this time, the bioreactor also reached a high cell density and

ready for use in kinetic and long-term stability studies.

5.2.6 Kinetic Studies

5.2.6.1 Well-Mixed

The bioreactor kinetics was first studied under continuous, well-mixed conditions for a period of -3 months (95 days). The bioreactor performance at various retention times up to 120 hrs were studied. At each feed rate, effluent sample was taken after at least two bioreactor volume feed had been collected from the bioreactor effluent to allow the bioreactor to reach pseudo-steady state. After each retention time study, the bioreactor condition was restored by feeding at ~ 2 0 hrs retention time to prevent any reactor upset caused by operation under extremely long retention time (> 50 hrs).

Samples were frozen immediately after pH measurement and stored for future analysis of

MAb, glucose, lactate, and glutamine concentration. 110 S.2.6.2 Plug-Flow

After completing the kinetic study under well-mixed conditions, the bioreactor was maintained at a fixed retention time of -20 hrs for about 40 days. The same reactor was then re-configured to a single-pass bioreactor. The bioreactor kinetics at plug-flow condition was then studied at various feed rates for a period of -35 days. The reactor was then changed back to well-mixed conditions for further study.

5.2.7 Long-Term Stability

The long-term stability of the fibrous bed bioreactor was tested under continuous, well-mixed conditions, at a fixed retention time (50 hours), a constant pH (7.0), and a constant dissolved oxygen level (30% saturation) for a period of 37 days. Samples were taken once per day and they were frozen in a freezer for future analysis after viable cell number had been counted.

5.2.8 Fed-Batch Kinetics

After completing the long-term stability study, the bioreactor was operated under fed-batch conditions. The continuous feed was turned off and the spent medium in the bioreactor was drained and replaced with fresh medium. After 48 hours, 5 ml sterilized,

PBS buffer solution containing 180 g/1 glucose and 160 raM glutamine was injected into the bioreactor to bring the glucose concentration back to 4.5 g/1 and glutamine to 4 raM.

The addition of glucose and glutamine was repeated again after another 48 hours.

Samples were taken from the bioreactor at 12-hour interval to measure pH, glucose, glutamine, lactate, and MAb. I l l 5.2.9 Analytical Methods

5.2.9.1 Cell Density and Viability

The total cell density was counted by diluting the sample in Dulbecco's phosphate

buffer saline (PBS) and counted on a hemocytometer. The cell viability was measured

with cell suspension diluted with 0.1% wt/v trypan blue in PBS and the viable cell

number was counted on a hemocytometer. The blue-colored cells were dead and

colorless, round, and bright were viable. The cell viability was estimated from cell

counts of viable and dead cells.

5.2.9.2 Glucose and Lactate Concentrations

The concentrations of glucose and lactate were measured using high performance

liquid chromatography (HPLC). The HPLC system consisted of a high pressure pump

(Waters, Model 6000A), an injector (Waters, U 6 K), an organic acid analysis column

(Bio-Rad, Model HPX-87H; condition: 45 °C), a column heater (Bio-Rad), a RI detector

(Waters, Series 410 differential refractometer; conditions: scale factor 20, sensitivity 16,

and internal oven temperature of 45 °C), and an integrator (Spectra Physics, Model 4270; conditions: chart speed 0.25 cm/min, attenuation 32; minimum area 5000, peak height 1,

and run time 20 minutes). The eluent was 0.01 N H 2SO4 at 0.6 ml/min flow rate. The details can be found elsewhere (Yang et al. 1992).

5.2.9.3 Monoclonal Antibody Concentration

The concentration of monoclonal antibody was analyzed using enzyme linked immuno-sorbent assay (ELISA) following the procedure described by Pierce. The details are given in Appendix B. 112 52.9A Glutamine Concentration

The glutamine concentration was determined using an enzyme assay kit from

Boehringer Mannheim (Cat. No. 139092), following the procedure described in the

manufacturer's instruction. Detailed procedures are given in Appendix C.

5.2.9.5 Cell Density in Fibrous Bed Bioreactor

At the end of this study, the immobilized cells in the fibrous bed bioreactor was

removed by washing the cells off the fibrous matrix with 0.5% trypsin in Dulbecco's

phosphate buffer saline (PBS) and counted with 0.1% (wt/v) trypan blue on a

hemocytometer. Five pieces of fibrous materials with size of 2x5 cm were cut from the

fibrous matrix. The weakly-bound cells were removed by dipping fibrous matrix 10

times in the buffer solution with known volume. The resulting cell suspension was then

counted for total cell number and viability. The fibrous matrix was then scrubbed hard to

wash off the remaining cells, which were counted as strongly-bound cells. The cell

number and viability of the freely suspended cells present in the bioreactor bulk fluid

were also counted.

5.2.9.6 Scanning Electron Microscopy (SEM)

After all the experiments had finished, fibrous matrix samples (lxl cm) were also

taken from different locations of the fibrous bed for examination of cell immobilization

condition under SEM. The SEM samples were prepared as follows:

After leaving the samples in 2.5% formaldehyde solution overnight, they were

rinsed 10 times with double distilled water, each time for 15 minutes. Then dehydrated

with 20 to 70% ethanol in the increment of 10% by leaving them in different concentration solutions for 30 minutes each time. Next day, they were dehydrated with

80% ethanol and twice with 95% and 100% ethanol for 30 minutes each time. The 113 samples were kept in 100% ethanol overnight again and then dried at critical point with liquid CO 2 by using a Pelco Critical Point drier in the Anatomy Department of the Ohio

State University. All steps, except that for the critical drying, were carried out at 4°C.

SEM photos were taken by using JOEL-Model 820 SEM at Geology Department of the

Ohio State University.

52.9.1 Confocal Microscopy

The fibrous matrix samples (lxl cm) taken from different locations of the bioreactor packing were fixed in 2.5% formaldehyde solution for 2 hours. Then, 40 Jig/1 propidium iodide (PI) solution (Sigma, No. P-4170) was used to stain the cells by immersing the fibrous matrix in the solution for 1.5 hours. Photos were then taken by using a Bio-Rad MRC 600 confocal microscope in the Anatomy Department of The Ohio

State University.

5.3 Results and Discussions

5.3.1 Effect of Cell Seeding Density

Six different seeding densities were tested to determine the optimal initial cell concentration. In general, regardless of the seeding density, all cultures had about the same lag period of about - 1 day and reached about the same maximum viable cell density of 1.6xl0 6 cell/ml (Figure 5.4). Except for the highest seeding density (6.4xl0 5 cells/ml), all cultures showed similar growth rate during the exponential phase, with an average doubling time of about 38 hours. In order to reach a high cell density faster,

3.2x103 cells/ml cell density was chosen for batch kinetic studies with static T-flask and spinner flask. 114 5.3.2 Kinetics of Batch Cultures of Hybridoma Cells

5.3.2.1 Static T-Flasks

The typical kinetics of batch culture of hybridoma cell HD-24 in a static T-flask is

shown in Figure 5.5. It took about 144 hours for the culture to reach a maximum cell

density of 1.60xl06 cells/ml and the MAb concentration of 43.64 mg/1. There was a long

lag phase of -36 hrs. The overall productivity of MAb was -0.3 mg I *1 h*1. As also

shown in Figure 5.5t the substrate glucose and glutamine were down to 0.114 g /1 and 0.1 1

mM, respectively, at culture time 144 hours. In the same time, lactate the growth

metabolite, increased to 2.89 g/1.

Figure 5.6 (a) shows the linear relationship between cell growth and glucose

consumption during the exponential phase, with cell yield of -3.9x10s per gram glucose

consumed. Figure 5.6(b) shows that the lactate yield from glucose during the entire

culturing period was constant, -0.64 g/g.

Figure 5.7 shows that the specific MAb productivity was almost proportional to

the specific growth rate. The specific MAb productivity was calculated by dividing the

derivatives found from the tangents on the curve of MAb production with the time by cell

number at that point. The specific cell growth rate was calculated using the non­

competitive growth equation (Equation 5 in Chapter 6 ). This indicates that hybridoma cells had higher productivity at faster growth rate, even though MAb is not necessary to

be growth-associated (Figure 5.8).

Figure 5.9 shows the linear plots of MAb concentration versus glucose concentration and glutamine concentration. In general, MAb production is not directly proportional to glucose and glutamine consumption, except for the exponential growth phase. 115 5.3.2.2 Spinner Flask

Figure 5.10 shows the hybridoma cell culture kinetics in a spinner flask. In general, the kinetics was similar to that found with T-flask culture. However, significant cell death occurred at -108 hrs, perhaps due to shear damage generated by the agitation in the spinner flask. The overall productivity of MAb was 0.30 mg l *1 h*1 and the maximum productivity was -0.44 mg l ' 1 h_1. The MAb concentration reached a maximum of 56.78 mg/1 at 192 hours. The glucose and glutamine concentrations were down to 0.121 g/1 and

0.17 mM, respectively, at 144 hours. Similar to the T-flask culture, lactate yield from glucose was constant (Figure 5.11), but MAb production was not always growth- associated and was not proportional to the substrate utilization (Figure 5.9).

5.3.3 Selection of Fibrous Materials for Cell Immobilization

The scanning electron micrographs of cell attachment conditions on various fibrous materials studied are shown in Figure 5.12. As can be seen in this figure, the

100% polyester yielded the highest cell density attached to the fibrous matrix. 100% cotton and 50% polyester/50% rayon also gave good cell attachment condition, but 100% glass fiber and 50% polyester/50% cotton had relatively poor attraction for the cells.

Thus, the 100% polyester fibrous material was chosen as the packing material for the fibrous bed bioreactor.

5.3.4 Kinetics of Hybridomas in Fibrous Bed Bioreactor

5.3.4.1 Well-Mixed Culture

Figure 5.13 shows the reactor kinetics under continuous well-mixed conditions.

The highest MAb concentration was 522.4 mg/1 obtained at the retention time 117.2 hours. The concentrations of glucose and glutamine decreased to zero at retention time

54.7 hours and 34.9 hours, respectively, indicating that the immobilized cells were under 116 starvation condition at the high retention time studied. Because of the depletion of

substrates, the MAb could not be produced at a higher level even a longer retention time

was allowed. The maximum reactor productivity was 7.4 mg I ' 1 h' 1 with MAb

concentration of 404.8 mg/1 at 54.7 hrs retention time. The reactor pH decreased from

7.2 to 6.7 due to the production of lactic acid.

5.3.4.2 Plug-Flow Culture

The kinetics for the fibrous bed bioreactor operated under single-pass conditions

are shown in Figure 5.14. Under plug-flow conditions, the reactor had a higher MAb

production rates at lower retention times as compared to those from well-mixed

bioreactor. The maximum reactor productivity was 16.4 mg I -1 h*1 at 134.9 mg/1 MAb

and 8.23 hrs retention time. At 45.26 hrs retention time, the MAb concentration was

447.84 mg/1 and reactor productivity was 9.89 mg l ' 1 h_1. However, the pH in the plug

flow fibrous bed bioreactor could not be adjusted. As the retention time increased, a pH

gradient developed along the length of the bioreactor due to the formation of lactic acid.

The eluent pH decreased from 7.2 to below 6.0 when the retention time increased to 45

hours or longer. The low pH in the bioreactor adversely affected the cell physiology and

inhibited MAb production, consequentially, the highest MAb concentration was only

443.73 mg/1 obtained at retention time 112.7 hours. .

As also shown in Figure 5.14, the MAb production at higher retention times might also be limited by the depletion of substrates. Glucose was almost depleted at 33.8 hours retention time and glutamine depletion occurred at retention time 24.7 hours.

5.3.4.3 Fed Batch Culture

The fibrous bed bioreactor was also operated under batch and fed-batch conditions and the results are shown in Figure 5.15. With continued additions of glucose and glutamine to the bioreactor, MAb production reached a maximum concentration of 117 668.9 mg/1. However, due to the accumulation of toxic metabolites and the low pH in the

bioreactor, the MAb productivity decreased dramatically at 108 hours. In general, the

productivity in the fed-batch bioreactor was lower than expected. This perhaps could be

attributed to that the bioreactor had been operated at a long retention time of 50 hrs for a

period of 36 days. The fed batch experiment was studied right after the reactor long-term

stability study. At that time, the reactor had not been recovered to its optimal condition.

Table 5.1 shows the MAb productivity and glucose and glutamine consumption rates

during the 3 cycles of the fed-batch culture. It is clear that the accumulation of toxic

metabolites and pH decrease inhibited the MAb production in the 3rd batch period, which

also can be seen by the dramatic decreases in the consumption rates of glucose and

glutamine. Figure 5.16 shows the lactate production from glucose and MAb production

from glucose and glutamine in the fibrous bed bioreactor under various operating

conditions. In general, the kinetics were similar to those in the suspended culture.

5.3.5 Long-Term Stability of the Fibrous Bed Bioreactor

Figure 5.18 shows the long-term stability of the fibrous bed bioreactor. As can be

seen in this figure, the concentrations of MAb, glucose, and glutamine in the bioreactor effluent remained almost unchanged during the entire 37-day continuous operation at 50 hrs retention time under well-mixed conditions. Compared with that in day one, the concentration of MAb in day 36 decreased only by 8.0%. The effluent cell density also remained at about the same level at all times. It is noted that before this experiment was conducted, the fibrous bed bioreactor had been operated for more than six months without suffering from any operation problems. The fibrous bed bioreactor has outstanding stability for long-term continuous operation. 118 5.3.6 Cell Density and Viability in Fibrous Bed Bioreactor

The total cell density and viable cell density in the bioreactor were found to be

1.466 xlO 8 and 1,17 xlO 8 cells/ml, respectively, at the end of this study. The majority of cells was immobilized in the fibrous matrix. Most immobilized cells were viable, but cells suspended in the bioreactor bulk fluid were most dead cells (Table 5.3). This indicated that the immobilization might have protected the cells from shear damage. It is also likely that dead cells were preferably washed out from the fibrous matrix, allowing the fibrous bed bioreactor to be renewed continuously and to maintain stable long-term production of MAb.

5.3.7 Comparison of MAb Production in Various Culture Systems

Figure 5.17 shows the comparison of MAb production in various culture systems.

It is clear that the fibrous bed bioreactor gave a much higher MAb production due to the high viable cell density attained in the fibrous bed, which was estimated to be -1.17x10s cells/ml at the end of this study. The productivity of MAb, maximum cell density, and highest MAb concentration in various culture systems are also compared and listed in

Table 5.2. It can be seen that the fibrous bed bioreactor had a MAb productivity of as high as 55-fold of that in the conventional suspended culture systems. The highest MAb concentration attained in the fed-Batch culture was 15.3-fold of that in T-flask culture.

The viable cell density in the fibrous bed bioreactor was -73 times of that in T-flasks.

However, the specific MAb productivity in the fibrous bed bioreactor was less than 40% of that in the static T-flask and spinner flask. This could be attributed to the high MAb and lactate concentrations in the fibrous bed bioreactor. It is known that high concentration lactate and ammonia are toxic or inhibitory to MAb production. The relatively low specific productivity in the well-mixed bioreactor also indicated the significance of product inhibition. Also, there might be significant nutrient limitation for 119 the immobilized cells in the fibrous matrix, which in turn would limit cell growth and

MAb production. As discussed early, specific MAb productivity seemed to increase with the specific growth rate of the hybridomas. Thus, the viable cells in the fibrous bed could not function as well as those in flask cultures because the limited availability of the substrates in the bioreactor. Also, there might be some mass transport limitations when there were high concentration of cell mass built up in the bioreactor. Mass transfer in the fibrous bed will be discussed in Chapter VI.

5.3.8 SEM Photos of Hybridoma Cells in Fibrous Bed Bioreactor

Figure 5.19 shows the morphologies of hybridoma cells immobilized in the fibrous matrix. It is clearly seen that high density of hybridoma cells were immobilized in the fibrous matrix by attachment to the fiber surface as well as by entrapment as large cell clumps within the interstitial spaces of the matrix. According to Needham et al.

(1991), five distinct cell morphologies can be identified for hybridoma cultures. (1) The small smooth spheres (some tethered) are the daughter cells in early G] phase of cell cycle; (2) Irregular shapes with small membrane blebs are the cells at late Gi or S phase;

(3) Large rough spheres are the cell at G2 phase; (4) Large smooth spheres with patterned interior and microvilliated are the cells at M phase; (5) The patterned, spherical cells undergoing cell division and they appear characteristic doublet that was also microvilliated are the cells at cytokinesis of M phase. The morphologies of hybridoma cells at various stages of the cell cycle can be found in these SEM photos, indicating that cells cultured in the fibrous bed bioreactor were in healthy condition and proliferating.

5.3.9 Confocal Microscopic Pictures of Hybridoma cell in Fibrous Bed Bioreactor

Figure 5.20 shows the confocal microscopic pictures of stained hybridoma cells located at different depth in the fiber matrix. There were relatively few cells at the surface of the fibrous matrix, perhaps due to the shear forces caused by fluid flow, that 120 had washed away cells from surface of the fibrous matrix. Slightly below the surface,

shear forces and mass transport resistance were small, and the matrix structure provided

an ideal environment for cells to grow and immobilize. Thus, a relatively large number

of cells were found. Further into the matrix, transport of substrates and toxic metabolites

became relatively difficult and in-efficient. Thus, the cell number decreased with

increasing the distance from the surface. However, the actual decline in cell number with depth was not as significant as seen in these photos, which only show stained cells. It was found that cells in the deeper parts of the fibrous matrix were not stained properly due to the insufficient time used in staining cell samples. Even at the center of the fibrous matrix, there still were large number of cells, but they were not stained and thus could not be seen in these photos.

5.4 Conclusions

The hybridoma cell HD-24 was successfully cultured in the fibrous bed bioreactor. The viable cell density reached 1.17x10s cells/ml in the bioreactor and the productivity of monoclonal antibody was about 55-fold as high as those in either spinner flasks or static T-flasks. The continuous operation of the bioreactor for 8 months and the long-term stability test showed that the productivity of the bioreactor was stable for continuous operation for a long period of time. Kinetic studies of substrate consumption, product formation, and metabolite generation showed that some physiological differences of cells cultured in various systems were due to availability of the substrate, removal of the metabolites, and culture conditions. The cell morphology analysis indicated that cells in the fibrous bed bioreactor were at different growth stages and they were healthy and proliferating. The confocal microscopy pictures gave a rough estimation of cell distribution at different depth in the fiber matrix. 121

Table 5.1 MAb productivity and glucose and glutamine consumption rates in the fed-batch culture

1st Batch 2 nd Batch 3 rd Batch (48 hrs) (48 h rs) (48 hrs)

MAb Concentration (mg/1) 293.74 642.52 668.92 F in a l pH 7.0 6.2 5.6 Lactate Concentration (g/1) 2.71 3.93 4.97 MAb Productivity (mg I'1 h'1) 6.12 7.27 0.35 Glucose Consum. Rate (g I '1!**1) 0.17 0.16 0.056 Glutamine Consum. Rate (m M /hr) 0.15 0.166 0.066

Table 5.2 Comparison of various systems for hybridoma cell HD-24 culture

Cutout MAb Max. MAb Max. Viable Specific Cell System Productivity Concentration Cell Density Productivity (m gr1 b '1) (mg/1) (cell/ml) (mg P1 ta^eeO-1)

Static T-Flask 0.30-0.60 43.64 1.60xl06 ~3.75xl0"7 Spinner Flask 0.300.44 56.78 1.57x10® ~3.82xl0‘7 Well-mixed Bioreactor -7.4 522.40 1.17x10® -6.32x10-® Plug-flow Bioreactor 9.9-16.4 461.23 1.17x10® “1.40x10 "7 Fed-Batch Bioreactor 6.0-13.5 668.92 1.17x10® “1.15xl0*7 122

Table 5.3 Cell viability and distribution in the fibrous bed bioreactor

CcUsin Weakly attached Strongly Attach bulk fluid cells cells

Total Cell Number (10®) S.2-6.5 195.0-208.0 52.2-60.0 Cell Distribution -2-.2.5% -75-80% -20-23% Viable Cell Number (10s) 0.06*0.1 136.6-166.4 46.4-54.0 Viability -1-1.5% -70*80% -89-90% Figure 5.1. The spirally wound structure of fibrous packing for the fibrous bed bioreactor 124

D.O./pH Meter

Sample Pori pH/D.O Probe r

»

Air/5% Medium 0 0 2 Recirculation Rask

Product

Figure 5.2 A schematic diagram of well-mixed fibrous bed bioreactor system D.O./pH Meter

Sample Port

Medium t H S B i # J ^ Product Tank spiner Air/5% C02 Flask

Figure 5.3. A schematic diagram of plug-flow fibrous bed bioreactor system Cell Density (celI/mL) 104 0 1 10 107 107 Figure 5.4 Effect of cell seeding concentration on cell growth cell on concentration seeding cell of Effect 5.4 Figure ® ® 1 3 5 7 6 5 4 3 2 1 0 ie (day) Time ------a — - ---- a— — ------— ♦ - - D ----- 2x10*4 8x10*4 4X10*4 3.2X10*5 6.4X10*5 1.6x10*5 126 Figure 5.5 Kinetics of MAb production by hybridoma cells cultured in static T-flask static in cultured cells hybridoma by production MAb of Kinetics 5.5 Figure

Glucose/Lactate (g/1) MAb (mg/l) 20 40 50 0 4 2 'Glutamine 3 1 5 0 0 0 * Glucose 30 060 30 ia (hr) Tima ia (hr) Tima 090 60 00 Lactate 120 ibe Celt Viable 120 A , MAb 150 150 5 I S e 3 E, >o JS 127 128

2.0

* 1 c • • ® O

0.5

0.0 0 1 2 3 4 Glucoa* (g/1)

(a)

3.0

y = 2.9190 - 0.64102* RA2 = 0.992

0.0 0 1 2 3 4 Glucott (g/l)

(b)

Figure 5.6 Linear plots of (a) cell density vs. glucose concentration and (b) lactate concentration vs. glucose concentration for T-flask culture 1.006-9 * y = 2.6994*01 + 4.3162e-8x RA2 = 0.975 > 8.006-10

Q. JX < Z _o 51 o 2 .006-10 • a (0 0.00e+0 — 0.000 0.005 0.010 0.015 0.020 Spoclflc Growth Rato (1/h)

Figure 5.7 Specific MAb productivity versus specific growth rate

Q> 30 E

0.0 0.5 1.0 1.5 2.0 Coll Donolty (10 * collo/ml)

Figure 5.8 Linear plot of MAb concentration versus cell density in the static T-flask culture Figure 5.9 Linear plots of (a) MAb vs. glucose concentration and (b) (b) and concentration glucose vs. MAb (a) of plots Linear 5.9 Figure MAb (mg/l) MAb (mg/l) 20 30 40 10 50 60 20 20 40 30 10 SO 60 0 0

1 3 5 4 3 2 1 0

1 3 5 4 3 2 1 0 e . ♦ MAb concentration versus glutamine concentration glutamine versus concentration MAb e 1 ■ ltmn (mM) Glutamln* lca (g/l) Glucoaa 1 ■ (b) (a) e Static Flasks Static e Flasks Spinner ■ d • Static Flask Static • 1 Spinner Flask Spinner

131

60 MAb

at 40 E o *

> e Viable Cell

0 50 100 150 200 Tim* (hr)

5

Glucose 3 4

Lactate ** A 5 o _i<8 •5 2 'Glutamine oI t o 2 1 (5

0 0 50 100 150 200 Tima (hr)

Figure 5.10 Kinetics of MAb production by hybridoma cells cultured in spinner flask Figure 5.11 Linear plot of lactate concentration versus glucose concentration glucose versus concentration lactate of plot Linear 5.11 Figure

Lactata (g/l) 0 2 1 3 0 1 lcs (g/l) Glucosa 2 3 ■ Static Flask Static ■ • Spinner Flask Spinner • 4

5 132 133

100% polyester

100% cotton 50% polyester/50% rayon

100% glassflber 50% polyester/50% cotton

Figure 5.12 SEM of hybridoma cells immobilized in various fibrous materials Figure 5.13 Kinetics of MAb production by hybridoma cells in fibrous bed bed fibrous in cells hybridoma by production MAb of Kinetics 5.13 Figure 200 S Glucose/Lactate (g/l) 400 300 500 600 100 0 4 <

0 f i 0

20 20 bioreactor under well-mixed conditions well-mixed under bioreactor aato Tm (hr) Tima Ratantlon aato Tm (hr) Tima Ratantlon 40 40 080 60 60 Lactate a Glutamine o Glucose ■ 80 100 100 MAb

120 120 <5 3 « E 134 Figure 5.14 Kinetics of MAb production by hybridoma cells in fibrous bed bioreactor bed fibrous in cells hybridoma by production MAb of Kinetics 5.14 Figure

Glucose/Lactate (g/l) 200 400 300 100 500 600 0 2 3 4 1 5 0 0 2 4 6 8 10 120 100 80 60 40 20 0 * 20 eeto Tm (hr) Time Retention eeto Tm (hr) Time Retention 060 40 under plug-flow conditions plug-flow under Glucose ■ Glutamine Lactate o * 80 100 MAb

—1 « 120 0 a 3 « E C 135 Figure S. IS Kinetics of MAb production by hybridoma cells in fibrous bed bioreactor bed fibrous in cells hybridoma by production MAb of S.Kinetics IS Figure

Glucose/Lactate (g/l) MAb (mg/l) 200 400 600 800 6 7 8 —i —r ' i i <— r i ■ "i— ' ■— i— i— r — 2 5 7 10 2 10 175 150 125 100 75 50 25 0 0 - , . 25 50 under fed-batch conditions fed-batch under ia (hr) Tima ia (hr) Tima 75 ------a— — .. o ---- 0 15 5 175 150 125 100 J < I I I I » < I J ' Glutamine Lactate Glucose MAb 0 o. 136 Figure 5.16 Kinetics of (a) lactate production from glucose, (b) MAb production from from production MAb (b) glucose, from production lactate (a) of Kinetics 5.16 Figure glucose, and (c) MAb production from glutamine in the fibrous bed bioreactor bed fibrous the in glutamine from production MAb (c) and glucose,

200 s (mg/l) Lactate (g/l) 600 400 500 300 100 3 0 2 1 w lca (g/l) Glucoaa (a) Batch ■ e Welt-Mixed ■ (b) Batch • Well-Mixed ■ Plug-Flow ♦ Plug-Flow

137 Figure 5.16 (continued) 5.16 Figure

MAb (mg/l) 2001 i 300 100 100 4001 4001 1 0 0 5 600 600 0 % 4 3 2 1 ltmn (mM) Glutamln* (c) _l_ • ■ ■ Welt-Mixed Plug-Row Batch - ■ - 138 Concentration 200 300 400 100 500 400 £ Figure S. 17 Comparison of MAb production in various culture systems culture various in production MAb of S. 17Comparison Figure O) 0 Figure S. 18 Long-term stability test of the fibrous bed bioreactor bed fibrous the of test stability S. 18Long-term Figure 0 200 100 500 600 700 0 20 10 40 ie (hr) Time ie (day) Time 60 iru Bd Bioreactor Bed Fibrous 20 0 0 10 140 120 100 80 Well-Mixed Spinner flask Spinner Static flask Static Fed-Batch Plug-flow a Glutamine (uM) Glutamine a (mg/l) MAb ■ o Glucose (mg/l) Glucose o 30

40 139 140

Figure 5.19 SEM photos of hybridoma cells in the Fibrous bed bioreactor 141

ZERO DEPTH 300 nm DEPTH

100 jim DEPTh 400 nm DEPTH

200 nm DEPTH S00 jim DEPTH

Figure 5 20 Confocal microscopic photos of stained hybridoma cells at different depth of Fibrous matrix 142

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2. Backer, M.P., L.S. Metzger, P.L. Slaber, K.L. Nevitt and G.B. Boder (1988). Large-scale production of monoclonal antibodies in suspension culture, Biotechnol. Bioeng., 32: 993-1000.

3. Broise D.D.L, M. Noiseux B. Massie, and R. Lemieux (1992). Hybridoma perfusion systems: a comparison study, Biotechnol. Bioeng., 40: 25-32.

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6 . Emery, A.N., M. Lavery, B. Williams and A. Handa (1987). Large scale hybridoma culture. Animal and Plant Cells Process Possibilities. Ellis Horwood Ltd., Chichester, 137-146.

7 Hagedom, J., and F. Kargi (1990). Coiled tube membrane bioreactor for cultivation of hybridoma cells producing monoclonal antibodies, Enzyme Microb. Technol., 12: 825-829.

8. Hulscher, M., CJ. Scheilber and U. Onken (1992), Selective recycle of viable animal cells by coupling of airlift reactor and cell settler, Biotechnol. Bioeng., 39: 442-446.

9. Katinger, H (1987). Principles of animal cell fermentation, Dev. Biol. Stand., 6 6 : 195-209 143 10. Klement, G., W. Scheier, and H. Katinger (1987). Construction of a large membrane reactor system with different compartment for cells, medium and products, Dev. Biol. Stand, 6 6 : 221-226.

11. Kohler, G., and C. Milstein, (1975). Continuous cultures of fused cells secreting antibody of predefined specificity. Nature, 256: 495-497.

12. Lambert, K.J., R. Boraston, P.W. Thompson and J.R. Birch (1987). Production of monoclonal antibodies using large scale culture. Development in Industrial Microbiology, 27:101-106.

13. Lee, G. M., A. Varma, and B.O. Palsson (1991). Production of monoclonal antibody using free-suspended and immobilized hybridoma cells : Effect of serum, Biotechnol. Bioeng., 38: 821-830.

14. Linardos T.I, N. Kalogerakis, L.A. Behie, and L.R. Lamontagne (1992). Monoclonal antibody production in dialyzed continuous suspension culture, Biotechnol. Bioeng., 39: 504-510.

15. Needham, H.P., Ting-Beall, and R. Tran-Son-Tay (1991). A physical characterization of GAP A3 hybridoma cells: morphology, geometry, and mechanical properties, Biotechnol. Bioeng., 38: 838-852.

16. Nilsson, K., W. Sheirer, O.W. Merton, L. Ostberg, E. Liehl, H.W.D Katinger and K. Mosbach (1983). Entrapment of animal cells for the production of monoclonal antibodies and other biomolecules. Nature, 302: 629-630.

17. Rupp, R.G. (1985). Use of cellular micro-encapsulation in large-scale production of monoclonal antibodies. Large -Scale Mammalian Cell Culture. Academic Press, Orlando, Pp 19-38.

18. Scheirer W. (1988). High-density growth of animal cells within cell retention fermentors quipped with membranes, In: Spier, R.E. and Griffiths, J.B. (eds.), Animal Cell Biotechnology, 3: 263-281. Academic Press Ltd., London.

19. Shi, Y., D.Y. Dewey, and H. Park (1992). Performance of mammalian cell culture bioreactor with a new impeller design, Biotechnol. Bioeng., 40: 260-270.

20. Sinacore, M.S. (1984). Gel entrapment: Applications in production of biological and mass culturing animal cells. Karyon Technology News I, Karyon Technology Inc., 333 Provident Highway, Norwood, MA 02062, U.S.A.. 144

21. Tharakan, J.P., and Chau, P.C. (1986). A radial flow hollow fiber bioreactor for the large-scale culture of mammalian cells, Biotechnol. Bioeng., 28, 329-342.

22. Wang, G.Z. W. Y. Zhang, C. Jacklin, D. Greedman, L. Eppstein, and A. Kadouri (1992). Modified CelliGen - packed bed bioreactors for hybridoma cell cultures, Cytotechnology, 4: 41-49.

23. Yang, S. T., I.C. Tang, and H. Zhu (1992a). A Novel Fermentation Process for Calcium Magnesium Acetate (CMA) Production from Cheese Whey, Appl. Biochem. Biotechnol., 34/35: 569-583.

24. Yang, S. T., H. Zhu, V. P. Lewis, and I. C. Tang (1992b). Calcium Magnesium Acetate (CMA) Production from Whey Permeate: Process and Economic Analysis. Resources, Conservation and Recycling, 7: 181-200.

25. Yang, S.T., H. Zhu, Y. Li, and G. Hong (1994). Continuous propionate production from whey permeate using a novel fibrous bed bioreactor. Biotechnol. Bioeng., 43: 1124-1130. 145

C H A P T E R V I HYBRIDOMA CELL KINETICS AND SIMULATION OF THE FIBROUS BED BIOREACTOR

6.1 Introduction

Hybridoma cells are cultured in different systems to produce a wide variety of monoclonal antibodies and have a quick expending market. Hybridoma cells, like other mammalian cells, grown in vitro exists in a dynamically complex and ill-defined environment (Glacken et al., 1988). Thus, understanding the roles of physiological and environmental factors on the growth and metabolism of hybridoma cells is prerequisite for the development of rational scale-up procedures (Miller, 1988).

Numerous models have been developed to describe the growth of hybridoma cells, the production of monoclonal antibodies, consumption of nutrients in the medium, and generation of metabolites. Bibila and Flickinger (1991, 1992) proposed a structured kinetic model to describe the correlation between cell growth phase and MAb synthesis and used this model to optimize the process of MAb production. Based on their results, environmental and/or genetic manipulation approaches could maximize the specific antibodies secretion rate and the antibody volumetric productivity in large-scale antibody production systems. Miller et al. (1988) and Ozturk et al. (1991) conducted independently kinetic analysis of the effects of nutrient, pH, and dilution rate on hybridoma growth and metabolism in batch and continuous suspension cultures. The rheological properties of hybridoma cell were studied by Shi et al. (1993). Dalili and

Ollis (1989), Heath (1989) and Lee (1991) investigated the effects of serum and other

145 146 energy sources on cell growth, metabolism, and antibody production. The kinetic

analysis by Glacken et al. (1988, 1989) suggested that the growth rate of hybridoma

could be described by Monod equation and glutamine had a noncompetitive inhibition on

cells, and that lactate was the only environmental parameter that significantly inhibited

antifibronectin MAb production by CRL-1606 hybridomas. Bree et al. (1988) studied the

relationship of glucose, glutamine, and lactate, and considered the rate of glucose consumption to be zero order in glucose concentration rather than the Monod model.

Ozturk et al. (1988) and Miller et al. (1988) concluded that the production of monoclonal antibodies was non-growth associated. However, in batch culture the period with varying specific growth rates is short and it is difficult to observe a wide range of specific growth rates (p) that are possible as with continuous cultures (Miller and Blanch, 1991).

Glacken et al (1988) showed that the specific MAb production rate was proportional to p when p was smaller than 0 .0 2 h r1 but independent of growth when p was greater than

0.02 h r 1. The specific production rate also depended on lactate concentration. Although there has been much literature published addressing the hybridoma cell growth and MAb production, most of them were based on the experimental data from static T-flasks and suspension culture.

Mathematical modeling of mass transfer phenomena and cell kinetics can indicate the dependence of nutrient, metabolite and product concentration inside the bioreactors

(Heath and Belfort, 1991). Numerous papers describing theoretical models of substrate diffusion/convection and uptake in different bioreactors have been published. Park and

Chang (1986) described the flow distribution in a hollow fiber bioreactor and the flow patterns were also visualized and modeled with magnetic resonance image techniques by

Heath et al. (1990). Models were developed for packed-bed and fibrous bed bioreactor bioreactors. Ethier (1981) investigated the creeping flow through mixed fibrous porous materials and found that vicious effects at the coarse fiber surfaces led to a significantly 147 lower overall permeability than that predicted by a simple application of Darcy’s law.

The pressure drop in a fixed bed of spherical particles was measured, calculated, and

predicted by Jaiswal et al. (1994). Avlontis (1993) and Boundinar (1992) improved the performance of spiral-wound (SPW) modules by optimizing some key geometrical parameters for given operating conditions. When the packed bed and fibrous bed bioreactors were employed to cultivate microorganisms and mammalian cells, many models coupling mass transfer equations with bioreaction kinetics were developed to simulate and optimize the bioreactors. Perry and Wang (1989) addressed the fundamental requirements of fibrous bed bioreactor for mammalian cell culture. Chiou et al. (1991) and Murakami (1991) modeled a concentric-cylinder airlift, fibrous bed reactor for mammalian cell culture and also studied its scale-up potential. The study of Park and

Stephanopoulos (1993) on a ceramic bead packed-bed cell culture bioreactor found that the extent of intraparticle convective medium flow was the dominant mechanism of nutrient transport to cells immobilized in the particle interior and the model thus developed was helpful for enhancing the limiting nutrients, such as oxygen, to allow maintenance of cell viability and productivity. The mass transfer and bioreaction kinetics in the biofilm on the surface of the packing material in the packed bed bioreactor was modeled by Skowlund and Kirmse (1989) and Lewandowski et al. (1991). Interactive dynamic programming (IDP) technique was used by Hartig and Keil (1993) for optimization of a large scale spherical fixed bed bioreactor and found the global optimum with higher probability compared with conventional methods. Despite their limitations, computer simulation can provide a significant savings of time, effort, and expense by allowing a quick and easy bioreactor inspection of concentration profile. Design, operation, and scale-up may also be significantly improved with the insights offered by the flexibility of theoretical investigations. 148 In this work, the fibrous bed bioreactor was used to cultivate hybridoma cell HD-

24. The kinetics parameters of cell metabolisms were estimated from the experimental

data from static flask culture. The governing equations of mass transfer in the fibrous bed

bioreactor were coupled with cell kinetics to develop a mathematics model and simulate the relationship between substrates, metabolites, and products in both static flasks and the fibrous bed bioreactor. The concentration profiles of substrates, metabolites, and products in the bioreactor were generated by the computer model. The fundamental design parameters were weighted and examined in terms of monoclonal antibody production and the effects of different design parameters were also simulated and studied.

6.2 Kinetics of Hybridoma Cell Culture

6.2.1 Kinetic Model

When hybridoma cells are cultured in static flasks under a batch mode, the rates of cell growth, substrate consumption and metabolite formation are generally expressed as following:

(1)

(2)

dP v dS m = ~r ' w (3) where t is time, ^ is the specific rate, Y \ is the cell yield, and Yp is the metabolite yield

The specific cell growth rate, \i, is usually dependent on substrate and product concentration. A double substrate-limited growth rate expression was proposed by Miller et al. (1988) for hybridoma, which incorporated ammonia and lactate inhibition. Growth was limited by glucose (S) and by glutamine(G): where S, G. A, and P are the concentrations of glucose, glutamine, ammonia, and lactate, respectively; it’max > K s, Kg , K a , and Kp are rate constants.

According to Glacken et al. (1988, 1989), lactate is the only environmental parameter that significantly inhibited antifibronectin MAb production by hybridomas. To simplify the equation, it is assumed that ( 1). compared to the inhibition effect of lactate, ammonia effect can be ignored, and ( 2). the concentration of glutamine is ample in the medium (G »K g) • Thus, Equation (4) is reduced to the following form:

(5)

After substituting the expression of it into equation (2), the following equation is obtained:

dS _ 1 dX _ 1 dP _ ~3t~YxHF YP~3t~ YX(KS +S)(Kr + P) (6)

The cell concentration (X) and lactate concentration (P) can be determined from glucose concentration (S) when cell yield, Yx, and lactate yield, Yp, are constant using the following equations:

X = Xo + Yx(S0 - S ) (7)

P = P0 + YP(S0 ~S) (8) where Xo, So, and Po are the initial concentrations of cell, glucose, and lactate, respectively. After substituting the expression for X and P into Equation ( 6 ) and performing integration, the following equation is obtained: 150 where (KP + P0- YpXq/Yx )(Ks +Sq+ Xq/Yx ) A = (10) So + Xq/Yx

Kt(KP + P0 +S0Yp) (11) So + X0fYx

The rate constants, Pmax, Ks, and Kp can be readily determined by using a least

squares analysis by comparing Equation (9) with batch culture data in the exponential

phase.

According to Bree et al. (1988) and Miller (1988), production of MAb from

hybridoma cells is partially growth-associated. A general model for MAb production is

as following: qAb =cqi -bp (12)

where q Ab is the specific productivity of MAb, p is the specific growth rate, and a and P

are constants.

6.2.2 Parameter Estimation

Equation (5) *(12) can be used to fit the data from T-flask experiments (Figure

5.5) and to find the best values for the parameters (pniax> Ks, Kp, Yx, Yp, a, and P) in the

model.

6.2.2.1 Growth and Product Yield

To determine the Yield factors, Yx and Yp, cell density, and lactate concentration

were plotted versus glucose concentration, respectively. As shown in Figures 6.2 and

6.3, both cell yield and lactate yield from glucose are constant in the exponential phase.

This also indicates that the lactate formation is growth-associated. Yield can be determined from the negative value of the slope. The cell yield, Yx, was found to be

3.89x10s cells/g glucose and lactate yield, Yp, was 0.641 g lactate/g glucose. 151 6.2.2.2 Specific MAb Productivity

The specific MAb productivity, qAb. was estimated from the MAb production rate divided by the cell density using the data shown in Figure 5.5. To determine the a and P

values in Equation (12), the calculated qAb data were plotted versus the calculated |i at

corresponding glucose and lactate concentrations using Equation (5) and the rate

constants determined later from model fitting. From this graph, a and P were estimated

as 4.316xl0-8 (mg/cell) and 2.699xl0*n (mg/h/cell), respectively.

6 .2.2.3 Rate Constants

Since Yx and Yp are constant, the integrated model. Equation (9), is valid to

describe the batch culture process of hybridoma ceils. The values of |imax. Ks- and Kp

can be estimated by fitting Equation (9) with concentrations of glucose at different

culture times using the software SIGMA-PLOT. The best values for those kinetic parameters were found as follows: Umax =0-92 h*1, Ks=2.579 g/l, and Kp=21.940 g/l. As

shown in Figures 6.5, 6 .6 , 6.7, the model prediction generally fit the data well. However,

at the early exponential phase, the cell densities were slightly higher than the data. This

is because that the model does not include ammonia inhibition effects, which might be

significant when lactate concentration was low. With the accumulation of lactate, the

inhibition effect from lactate gradually became dominate and the ammonia effect became insignificant. All the kinetic parameters thus obtained are listed in Table 6.1.

6.3 Mathematical Model for Fibrous Bed Bioreactor

Mass transfer models can provide the necessary theoretical background in the request for optimum use of the fibrous-bed bioreactor for growing animal cells and harvesting their products. However, rules governing the design and operation of the 152 structured fibrous bed for cell culture applications do not exist in the literature. Instead,

system development is generally based on experiment and accumulated experience.

6.3.1 Pressure Drop and Shear Force in the Fibrous Bed

According to Perry and Wang (1989), the pressure drop in a fibrous bed can be

calculated by considering the fibrous bed as a porous medium. Brinkman's equation is

applied:

VP = TjV2u —^j~u (13)

where P is pressure, u is local fluid velocity, r| is viscosity, and k is permeability, defined

so as to account for the force exerted by the neighboring fibers on the fluid. By using

Darcy's Law,

< i 4 >

where L is the reactor length and U is the fluid velocity far from a fiber.

The shear force exerted on cells in the fibrous bed can be determined by solving

the relevant hydrodynamic equations by Spielman and Goren (1968).

FJ L = n U _d_ s k 4(1- e ) where Fd is drag force, s is surface area, d is the diameter of fiber, and e is the porosity of the bed. The permeability k can be evaluated iteratively (Koch, 1986) by using the following equations and the software MAPLE:

r = ^ f < l6> where , _ r K,(r/k1‘) . I r ’ , /. - -gfflprj + 2T1 (l7) 153 and (18) where r is the radius of fiber and Ko(x), K i(x) are modified Bessel functions of order zero and one, respectively. In this work, the length of the fibrous bed, L, is 10 cm, the medium viscosity q is

0.7 cp at 37 °C, and fiber radius, r, is 12 pm. By using Equation (14), the pressure drop as a function of the liquid velocity, U, was calculated and is shown in Figure 6 .8.

The permeability, k, was determined using Equation (16) and was found

2.063x1 O'10. The shear force on cells immobilized in the fibrous bed can be evaluated by using equation (15). With the following conditions: fiber diameter, d=24 pm, porosity, e

= 66.3%, and liquid velocity, U=l. 1 lxlO ”6 m/s, the shear in the fibrous matrix, Fd/s, was found to be 8.478xl0'3 N/m2, which is much less than the level that could cause significant cell damage, according to the research conducted by Petersen et al. (1988) and

Bom etal. (1992).

6.3.2 Mass Transfer Coefficient in the Fibrous Bed

Transport of nutrients to cells at the fiber-medium interface is analogous to particle collection by the diffusive mechanism in fibrous filter (Perry and Wang, 1989).

The collection rate per unit volume (f) is (Spielman and Goren, 1968):

r = 0.9112^- 9$-vPe-*3) (19) where n is number of particles per unit volume, Pe is Peclet number and equals to Ud/D, D is molecular diffusivity, and ^ is a flow parameter given by:

c _ KoirfkV) W"- • ) (20) 5 “ (rfkv2)Kt(r/kvl)

The collection rate corresponds to the mass transfer coefficient, kc, as follows: 154

kc = 0.9lM (2.0$-*3 Pe-*3) (21) it

Glucose was the limiting substrate. The mass transfer coefficient for glucose in

the fibrous bed was estimated by using equation ( 21) under the following conditions: length of the fibrous bed, L=10 cm, the medium viscosity r|=0.7 cp at 37 °C, fiber radius

r=12 ^m, bed porosity, £=86.37%, and diffusivity of glucose in water, D *=6.73x10 ‘ 10

m2/s (Lide, 1991). The mass transfer coefficient, kc, was thus calculated as 1.375xlO -8

m/s when U =l.l lxlO' 6 m/s.

6.3.3 Mass Transfer with Reaction in Fibrous Bed

The general mass transfer equation with chemical reaction is as follows:

%± + {U*VS) = Dt V2S + R (22)

where S(x,y,z) is the concentration of substrate (usually the limiting one), t is time, U is the velocity vector (w, u, v), and D e is the effective diffusivity of substrate.

The reaction rate (R) is dependent on the specific growth rate (ti) and cell concentration

(X), as follows:

(23) where is a function of S, and P following in Equation (5).

The spirally-wound fibrous-bed bioreactor has a structured packing (Figure 5.1).

The side view of the fibrous matrix packing is shown in Figure 6.1. In the fibrous-bed bioreactor, liquid medium is flowing in axial direction only. To simplify the model equation, the following assumptions are made: ( 1). homogeneous cell suspension in the cell region, (2). an isothermal system, (3). Fickian diffusion characterized by an effective 155 diffusivity D e, and (4). steady-state substrate conversion. With these assumptions,

Equation (22) in Cartesian coordinate system reduces to: U-§=D'<0 + I£) + * (24> where Um is the liquid flow velocity in the matrix phase.

The diffusion term in the z-direction in Equation (24) can be neglected. Equation

(24) becomes:

u"% = D‘l£ + R (2 5 )

The boundary conditions for Equation (25) are:

J | = 0 at x = 0 (26)

atx = a (27)

S = S0 at z = 0 (28) where Sg is the substrate concentration in the gap phase. Since the cells in the gap phase was only about 2% of the total cells in the bioreactor, reaction in the gap phase is negligible and thus, we assumed that Sg = So = 3.52 (g/1).

6.3.4 Liquid Flow Velocity

In the gap phase, the linear flow velocity can be determined by using the following equation: 156 where L is the total length of the fibrous matrix, T| is the viscosity of the medium, and Ug,

Qg, Ag, AP, and b are the linear velocity, volumetric flow rate, cross-sectional area, pressure drop, and half of the width in the gap phase, respectively.

In the fiber matrix phase, the pressure drop can be described by Darcy's Law:

U — f i w — ~ A P k Um An L T) tJU) where L is the length of the fibrous bed, r\ is the viscosity of the medium, k is permeability, and Um, Qm, Am, and AP are the linear velocity, volumetric flow rate, cross-sectional area, and the pressure drop in the fiber matrix phase, respectively.

Since the pressure drop in both phases are equal, Equations (29) and (30) can be combined as follows:

U m _ 3k (31) U s ~ W or

6.4 Simulation of Fibrous Bed Bioreactor

Equation (25) can be solved numerically to find the concentration gradient in the fibrous bed bioreactor. A FORTRAN code (Appendix A) was written to link the software IMSL to solve the partial differential equation. Unless otherwise mentioned, parameter values listed in Table 6 .1 and Table 6.2 were used in computer simulation study of the fibrous bed bioreactor.

6.4.1 Concentration Gradients in Fibrous Bed

The concentrations of MAb, glucose, and lactate in a plug-flow fibrous bed bioreactor can be generated by using equations (5), (7), ( 8), (25), (26), and (27). The 157 concentrations of MAb, glucose, and lactate in a plug-flow fibrous bed bioreactor with a

retention time of 15.4 hrs are shown in Figures 6.9, 6 .10, and 6 .11, respectively.

The concentrations of glucose, lactate, and MAb change along both axial and

radial directions. In the axial or bioreactor length direction, MAb and lactate

concentrations increase while glucose concentration decreases. In the radial or fiber

matrix thickness direction, there were dramatic concentration changes near the surface of

the fibrous matrix. However, the concentration gradient became flat after a short distance

from the surface, suggesting that there might be diffusion limitations between the fibrous

matrix phase and the bulk liquid phase when the linear flow velocity is small. However,

when the retention time decreases or linear flow velocity increases, convection in the

axial direction in the fibrous matrix is enhanced and the concentration difference at the

interface minimizes.

6.4.2 Comparison of Results from Experiments and Model Simulation

Figure 6.12 shows the outlet concentrations of MAb, glucose, and lactate

generated from model simulation and compares them with the experimental data from the

plug flow fibrous bed bioreactor (Figure 5.14). The model simulation was under ideal

and theoretical conditions, which was not the same as the actual experimental situations.

The viable cell density in the bioreactor was adjusted in the simulation to allow good fit

with the experimental data. The simulated concentration profiles shown in Figure 6.12 are based on a viable cell density of 6.67xl06 cells/ml. Although this number is much lower than the actual cell density found in the bioreactor, the shape and trend of the simulation curves thus generated are very close to the experimental data.

The model was thus used to study several important reactor parameters and their effects on MAb production. 158 6.4.3 Effect of Porosity of the Fibrous Matrix on MAb Production

The porosity of fibrous bed without cells was -0.925. The bed porosity is a

function of immobilized cell density. When cell density in the bioreactor reached

2.11x10s cells/ml, the porosity of the fiber bed was calculated to be 0.45 by assuming that the cell had an average diameter of 15 p.m and were evenly distributed in the fiber bed.

The porosity affects the mass transfer coefficient and linear flow velocity of medium in

the matrix, and thus the productivity of MAb.

The bed porosity is a linear function of cell density in the bioreactor and affects the mass transfer coefficient and linear flow velocity of medium in the matrix and in the gap. In the simulation, a range of fibrous bed porosity from 0.45 to 0.925 was studied.

The liquid medium flow velocity in the fibrous matrix can be calculated by using equations (30) and (32). Figure 6.13 shows that when the porosity of the fibrous bed bioreactor decreases, the MAb productivity increases dramatically because the cell density in the bioreactor increases. However, when the porosity reaches 0.5, the MAb productivity starts to decrease due to mass transfer limitations caused by large cell mass present in the fibrous matrix that reduces the medium flow in the fibrous matrix. As shown in Figure 6.14, the ratio of the volumetric flow rate in the matrix phase to that in the gap phase (Qm^Qg) decreases with a decrease in the porosity or increase in cell density. The specific MAb productivity also decreases with decreasing bed porosity due to the increased diffusion limitation.

6.4.4 Effect of the Thickness of Fiber Matrix

Different matrix thickness changes the length of mass transport path of substrate diffusion, thus affecting the substrate delivery, metabolite removal, and MAb production.

In this simulation, the diffusion effect in the radial direction at different matrix thickness 159 was examined by varying fiber matrix thickness from 0.1 to 100 mm, while keeping the thickness ration constant at a/b = 1.

When the volumetric flow rate was kept constant, the effect of different thickness of matrix on MAb production is shown in Figure 6.15. Different matrix thickness changes the length of mass transport path of substrate diffusion, thus affecting the substrate delivery and metabolite removal. In this simulation, when the thickness of the fibrous matrix increases, the diffusion path in the matrix phase increases, which means that the mass transfer limitation becomes more severe, so that the MAb production decreases as the thickness of the fibrous matrix increases.

6.4.5 Effect of the Thickness Ratio of Matrix to Gap, a/b

Changing the thickness ratio of matrix phase to gap phase would affect the overall cell density in the bioreactor, linear flow velocity, and volumetric flow rate in both phases. The effect of different thickness ratio (from 0.1 to 10) on MAb production in the fibrous bed bioreactor was thus studied for a fixed a = 2mm.

Figure 6.16 shows the effect of the thickness ratio of matrix phase to the gap phase on MAb production. The change of thickness ratio affects the total cell number in the fibrous bed bioreactor and volumetric flow rate in both phases. When the ratio of the thickness of matrix phase to gap phase is small (0 . 1- 1.0 ), change of the ratio affected the

MAb production greatly. When the ratio is higher than 1.0, MAb production by the ratio change is not as sensitive as those in the small ratio range. This is because when the thickness ratio increases, the convection in the matrix phase increases and the diffusion in the radial direction became less significant. Figure 6.17 shows that when the thickness ratio increases, the ratio of volumetric flow rate in matrix and gap phase increases and the specific MAb production decreases greatly when the ratio is small (0.1-1). However, 160 when the ratio becomes larger, the specific MAb production decreases slowly and even has a slight increase when the ratio was larger than 4.

6.4.6 Effect of Recirculation Rate on MAb Production

When a recirculation was implemented into a plug-flow bioreactor system, the recirculation rate determined the mixing condition in both matrix and gap phases, and the mode of bioreactor operation. The effect of recirculation rate on MAb production was studied at a fixed a retention time of 10 hours.

Figure 6.18 shows the effect of recirculation rate on MAb production in the fibrous bed. Increase in the recirculation rate would increase the convection flow rate in the matrix phase, which results in an increase in the MAb production because more substrates were available to the immobilized cells. However, further increasing recirculation rate would change the bioreactor from plug-flow to well-mixed conditions, and reduce the bioreactor efficiency and MAb productivity. This simulation study showed that the optimal recirculation ratio was 10 at a retention time of 10 hours in the fibrous bed bioreactor. 161

Table 6 .1 Parameters and their values for the kinetic model for hybridoma cell culture

P aram eter Value

Ks 2.58 g/1 KP 21.94 g/1 Umax 0.92 h-1 Y x 3.89x10s cell/g glucose Yp 0.614 g lactate/g glucose a 4.32xl0*8 mg/cell P / 2.7QX10"11 mg h' 1 cell' 1

Table 6.2 The parameters and their values used in computer simulation of the fibrous bed bioreactor

Param eter Value

Fibrous bed length (L) 10 cm Fiber radius (r) 12 Jim Half width of matrix (a) 2 mm Half width of gap (b) 4 mm Fibrous bed porosity (£) 86.37% Medium viscosity (r|) 0.7 cp Effective diffusivity (D e) 6.73xl0 *10 m2/s Mass transfer coefficient (kc) 1.38x10-8 m/s Center line Center line of matrix of gap a I

D i f f u 3

MMH taP Convection

Figure 6.1 Schematic Diagram of Mass Transfer regime in Fibrous-Bed Figure 6.3 The linear plot of lactate concentration versus glucose concentration glucose versus concentration lactate of plot linear The 6.3 Figure Cell Density (cell/ml) Figure 6.2 The linear plot of cell density versus glucose concentration glucose versus density cell of plot linear The 6.2 Figure 0.0 2.0 3.0 1.0 0 0 y = 1.6430e+6 * 3.8893e+5x RA2 = 0.966 = RA2 *3.8893e+5x 1.6430e+6 = y 1 299 -0612 RA20.992 = 0.64102* - 2.9190 = y 1 lcs (g/1) Glucose lcs (g/1) Glucose 2 2 4 3 3 4 163 £ y = 2.6994e-l 1 + 4.3162e-8x RA2 = 0.975 > 8 .006-10

I E.

o o a. CO 0.00e+0 — 0.000 0 .0 0 5 0.010 0 .0 1 5 0.020 Specific Growth Rato (1/h)

Figure 6.4 The relationship between specific MAb productivity and specific growth rate

17 16 * 15 14 13 o 12 10 11 o 10 x 9 8 7 • Experiment m 6 C v Predicted Q) 5 Q 4 3 h 2 1 0 J I ■ I ■ I i I ■ I i I i I . I . I . L 30 40 50 60 70 80 90 100 110 120 130 140 150 Time (hour)

Figure 6.5 Hybridoma cell density change during static flask culture 165

5.0

4.5

4.0

£ ? 3.5 • Experiment ^29 3.0 V P red icted

mW 2.59 S o V • g 2.0 V • v • 3 1.5

1.0

0.5 0 7 0.0 -J * 1 * I J I l 1 1 I 1___- L * - 1 L I ■ I t i 30 40 50 60 70 00 90 100 110 120 130 140 150 Time (hour)

Figure 6.6 Glucose concentration change during static flask culture

4.0 t— i— i—■—i— ■— r - i ■ i J i—* - 1 ->— i—•—r ■ » -

3.5 - -

3.0 - _ 0 V . 0 7 \ 2.5 _ - taO m0V0V ■ - V 2.0 - y - V 1.5 _- 3 v . # • E xperim ent 1.0 • V P redicted 1 • 0.5 -

0.0 ■ , i . i . i . i—*—i—.— . 1 __i__i__i__i—i—.— 30 40 50 60 70 80 90 100 110 120 130 140 150 Time (hour)

Figure 6.7 Lactate concentration change during static flask culture Figure Figure Pressure Drop (N/mA2) Figure 6.9 The 3-D MAb concentration profile in the fiber matrix fiber the in profile concentration MAb 3-D The 6.9 Figure 0 0 . 2 4 3 1 5 00+0 — ■ 6.8 Relationship between pressure drop and medium velocity medium and drop pressure between Relationship

00-5 0 .0 5 li Vlct (m/s) Velocity Fluid in the fibrous matrix fibrous the in 50-4 0 .5 1 00-4 0 .0 2 [■MM 00[■MM ■o-so □ 100-150 100-150 □ D D 150-200 ■ 0 0 -3 0 S 2 ■ 20 O- 2 S 0

167

U* » 2i W ft © „ O ^ ° ©

Figure 6 .10 The 3-D Glucose concentration profile in the fiber matrix

Figure 6.11 The 3-D Lactate concentration profile in the fiber matrix Figure 6.12 Comparison of experimental data and model simulation of of simulation model and data experimental of Comparison 6.12 Figure M Ab(1 OOmg/l) 0 2 3 4 1 5 0 hybridoma culture in plug-flow fibrous bed bioreactor bed fibrous plug-flow in culture hybridoma 0 70 0 6 0 5 0 4 0 3 0 2 10 eeto Tm (hr) Time Retention MAb Lactate lcs ' Glucose 5 • * 3 o o m U - Simulation -- Experimental •

MAb (mg/2x10 Cells) MAb (mg/l) Figure 6.14 Effects of bed porosity on volumetric flow rate ratio ratio rate flow volumetric on porosity bed of Effects 6.14 Figure 200 100 100 150 50 0 2 4 3 5 1 0 4 0.5 0 .4 0 4 0. 6 0. 8 0. 1.0 .9 0 .8 0 .7 0 .6 0 .5 0 .4 0

Figure 6.13 Effects of bed porosity on MAb production production MAb on porosity bed of Effects 6.13 Figure Qm/Qg MAb 6 0. .8 0 .7 0 .6 0 and specific MAb production MAb specific and and glucose consumption glucose and Glucose Porosity Porosity 9 1.0 .9 0 « 3 200 300 100 (9 o o o E 169 Concentration s < A 190 E Q» 1 2 3 4 5 60 50 40 30 20 10 0 IV Figure 6.16 Effect of the recirculation ratio on MAb production MAb on ratio recirculation the of Effect 6.16 Figure 210 160 200 170 6 0.4 160 0 Figure 6.15 Effects of thickness of fiber matrix on MAb MAb on matrix fiber of thickness of Effects 6.15 Figure eiclto Ratio Recirculation Glucose MAb 20 ie Tikes (mm) Thickness Fiber production and glucose concentration glucose and production 0 6 0 4 60 100 o Glucose (g/1) Glucose o a Lactate (g/1) Lactate a MAb n 0.2 0.0 0.6 0.6 1.0 (xO.Ig/l)

170 MAb (mg/2x10 Cells) MAb (mg/l) on specific MAb production and the volumetric flow rate ratio rate flow volumetric the and production MAb specific on 120 160 40 Figure 6.18 Effect of the thickness ratio of matrix to gap gap to matrix of ratio thickness the of Effect 6.18 Figure Figure 6.17 Effect of the thickness ratio of matrix to gap gap to matrix of ratio thickness the of Effect 6.17 Figure 0 on MAb production and glucose consumption glucose and production MAb on 4 8 6 4 2 2 hcns Rto (a/b) Ratio Thickness hcns Ratio Thickness 4 Glucoss Ab M 6 Qm/Qg B 1 0.0 0.1 0.2 0.4 0.5 0.3

0.6 0 0 0 0 o O «• E o» 171 172

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A computer program for simulation of the fibrous bed bioreator

186 c c c c c nno onn non non non nnonnnonnnnnnn non c DIFINITION OF VARIABLES OF DIFINITION HRCE TITLE*30 CHARACTER USRT OCNRTO (g/1) CONCENTRATION SUBSTRATE - Y (g/l/a) RATE REACTION (m/a*c) NX COEFFICIENT (kg/m*3) TRANSFER MASS - CONSTANT KX (kg/m*3) INHIBITION * KP CONSTANT REACTION * KM DM ■ DIFFUSION COEFFICIENT FOR REACTANT (m*2/B*c) (m*2/B*c) REACTANT FOR (kg/m*3) COEFFICIENT CONCENTRATION DIFFUSION ■ DM REACTANT INITIAL - CO NEE LY NDS NX NPDES, LDY, INTEGER NRNI FOT SQRT FLOAT, INTRINSIC NSTEP NOUT, IDO, K, J, I, INTEGER A COMPUTER SIMULATION OP MASS TRANSFER AND KINETIC KINETIC AND TRANSFER MASS OP SIMULATION COMPUTER A IN A CONTINUOUS FIBROUS BED BIOREACTOR BIOREACTOR BED FIBROUS CONTINUOUS A IN ERAY 1995 FEBRUARY, HUI ZHU ZHU HUI EXTERNAL FCNBC, PCNUT REAL FCNBC, FCNUT OPEN (UNIT*30, FILE-'FINAL.out', STATUS-'NEW)

C SET BREAKPOINT fc INITIAL CONDITIONS *

DO 10 1*1, NX XBREAK(I)* (FLOAT!I-1)/FLOAT(NX-1))

C SET FUNCTION VALUES

Y(l, I) *3.35 10 CONTINUE

C SET PARAMETER FOR MOLCH

TOL*SQRT(AHACH(4)) HINIT-0.01*TOL T*0 .0 IDO-1 NSTEP-100 CALL UMACH (2, NOUT) J— 1 20 CONTINUE J*J+1 TEND-FLOATtJ)/FLOAT(NSTEP)

C SOLVE THE PROBLEM •

CALL MOLCH (IDO, FCNUT, FCNBC, NPDES, T, TEND, NX, XBREAX, * TOL, HINIT, Y. LDY)

C PRINT RESULTS * C********************************************* IF (MOD(J,10).LT.0.000001) THEN WRITE (30, 100) (Y(LDY, K), K*l, NX) 100 FORMAT (IX, 10F6.3) END IF WRITE (TITLE,'(A, F4.2)') 'SOLUTION AT T*', T CALL MRRRN (TITLE, NPDES, NX, Y, LDY, 0)

C FINAL CALL TO RELEASE WORKSPACE *

C IF (J.BQ.5) QOTO 30 IF (J.BQ.NSTEP) IDO-3 IP (J.LB.NSTEP) OOTO 20 non non non n o n CONTINUE 3 0 (KP+2.919-0■841*0(1))/YXS+M)/UM • KX-1.37SE-8 KP-21.94 2.579KM- DEFINE DEFINE BOUNDARYTHE CONDITIONS REAL X, T,ALPHA(l), BETA(1), OAMP(l),CM KX. SPECIFICATIONSFOR ARGUMENT SUBROUTINEFCNBC(NPDES, X, T, ALPHA,BETA, GAMP) UT (I)--DH*UXX(1)/UM-C*(VMAX*U<1)/(KM*U(1))/ M-2.C04*-14 UM-1,l79*-5 VMAX-2.56E-4 D-6.73E-10 YXS-3.889X5 END RETURN EMDIF ELSE DK-S.734-10 END RETURN C-6.67EC DEFINE FDETHE UXX(l), EM, UT(1), VMAX, X. T,U(l), UXU), REAL URUIEFNT (NPDES,SUBROUTINE FCNUT X, T,U,UX, UT) UXX, F (X ,EQ.IF 0.0) THEN INTEGER NPDES PCFCTO O RUET * SPECIFICATION FOR ARGUMENTS END NEE NPDES INTEGER GAMP(1)-0.0 SAMP (1)>0.0 ALPHA(l)-0.0 ALPHA(1)-1.0 1ETA(1)-(DM/XX) BETA(11-1.0

KM, KP, YXS,UN, M,C

189 APPENDIX B.

ELISA for MAb determination

190 1. STOCK REAGENTS AND STORAGE:

- 0.2 M Sodium-Bicarbonate (Coating Buffer, pH 9.4). Store 4 °C.

- 10X Phosphated-Citrate Buffer (pH 5.0)

Dissolve 4.67g Citrate and 7.30g Sodium Phopht. (Dibasic) in 150 ml

ddH 2 0 . Adjust pH. Store 4°C.

- 0.1 M PBS Solution (pH 7.4): Dissolve 8.0g NaCl, 1.16g Na 2HP0 4 , 0.20

KH2PO4,0.20g KC1 in 1 0L ddH 20. Adjust pH.

- Blocking Solution: 0.1M PBS Solution (pH 7.4) w/10% horse serum. Store

4°C.

- PBS/Tween20 0.05%: Add 0.5 ml Tween to 1L of PBS Solution. Store 4°C.

- 30% H 2O2. Aliquated and store -20°C.

- Stop Solution: 0.5-2.0 M H 2SO4.

- DMSO (dimethyl sulfoxide)

- TMB, Free Bae tablets (Pierce# 34016X);

- Rabbit anti-Mouse 1 °Ab (Pierce # 31188), aliquot and stored at 4°C.

- Rabbit anti-Mouse 2°AB (Pierce #31450X), aliquated and stored at 4°C.

- Standard: mouse IgG2b (Sigma #M7644)

D. PROCEDURE

C oating

1. Dilute 1 °Ab to 5 p.g/ml in Coating Buffer.

2. Place 100 til diluted 1 °Ab in each well.

3. Cover and incubate for 6 hr at 23°C (or 18 hrs at 4°C)

4. Wash 2x100 p.1 with PBS/Tween buffer.

B locking

1. Place 100 ill standards and samples per well. 192 2. Cover and incubate for 1 hr at 37 °C.

3. Wash 4x100 )il with PBS/Tween buffer.

2°A b

1. Diluted 2°Ab to 0.1 tig/ml in PBS/Tween.

2. Place 100 p.1 dilute 2°Ab in each well.

3. Cover and incubate for 1 hr at 27 °C.

4. Wash 6x100 ill with PBS/Tween buffer.

Detection

1. Dissolve one TMB free base tablet in 1 ml DMSO. Add this solution to

8 ml ddH 2 0 w/mixing. Add 1 ml of lOxPhosphate-

Citrate immediately prior to use add 10 ill30% H 2O2.

2. Place 100 ill dilute TMB in each well.

3. Cover and incubate for 30 min at 23°C.

4. Stop reaction with 100 ill 0.5-2.0 M H 2SO4.

5. Read Absorbance at 450 nm. III STANDARD CURVE STANDARD III

Concentration (mg/L) 0 1 * 0 4 5 3 2 0.0 1 Figure B.l Standard curve for MAb determination MAb for curve Standard B.l Figure 20066 10A(4.9019x) * 2.00016*6 = 0.2 .4 0 0.6 0.8 1.0 1.2 APPENDIX C

Glutamine Analysis

194 195 I. STOCK REAGENT

- Glutamic acid Kit from Boehringer Mannheim (Cat.# 139092).

- L-Asparaginase from Boehringer Mannheim (Cat.# 102903).

D. PROCEDURE

Wavelength: (hg) 492nm

Glass cuvette: 1 cm light path

Temperature: 20-25 °C

Final volume: 3.03 ml

1. Addition of Reagents and Sample into a Cuvette as Table C.l

Table C. 1 The procedure of reagents and sample addition

Pipette into cuvettes Blank (ml) Sample (ml)

Solution 1 0.6 0.6 Solution 2 0 .2 0 .2 Solution 3 0.2 0 .2 Redist. water 2.0 1.8 Sample solution - 0 .2

2. Mix, after 2 min read absorbances of the solution (Ai). Repeated the measurement after 2 min until the difference less than 0 .0 10 /2 min).

3. In case of constant absorbance, start the reaction by addition of 0.03 ml Solution 4 into blank and sample, respectively. Mix, wait until the reaction has stopped (approx. 15 min) and read absorbance of the blank and sample immediately one after another (A 2). 196 4. Add 0.06 ml L asparaginase into blank and sample cuvettes, respectively. Mix, wait

until the reaction has stopped (approx. 15 min) and read absorbance of the blank and

sample immediately one after another (A 3).

m . CALCULATION

A A b -g lu ta m m , A A talnpifL-tluta min t ~ ^^btmk-thuanine ( 0

VxMW , n c ExdxvxlOOx ^ J

where V=final volume(ml); v=sample volume (ml); MW=molecular weight of glutamine; d=light path (cm); and e=absoiption coefficient of formazan at 492 nmsl9.9.

c - W i ^ M x / o O O ^ ■ 0 .1 I3 S X M U H 1 (3)

IV. STANDARD CURVE

80 ■ y = 9.1835 + 169.52X RA2 = 0.994 70

60

50

40

30

20 0.1 0.2 0.3 0.4 Dalta(A)

Figure Cl. The Standdard curve for glutamine measurement APPPENDIX D

Cell Density Determination

197 198 I. PROCUDURE

The cell density in the fibrous bed bioreactor was determined from the total protein content in the sample. In general, the total protein content is proportional to the cell number or density. The total protein content was measured by dye-binding protein assay (Bradford method)6. Bovine serum albumin (BSA) was used as the standard for the protein assay. Five small pieces of fibrous matrix (2.0x2.0 cm) were cut from the bioreactor packing matrix and used for protein assay and cell density estimation.

Samples and a series of standard BSA solutions (0-100fil, mg/1) were prepared and mixed with appropriate amount of NaOH (IN) to a final volume of 200 ill. Then, 5 ml

Coomassie Blue G-250 (Sigma) were added into the tubes and the tubes were votexed.

The standards and samples were read in a Beckman DU 640 spectrophotometer. The standard curve of protein assay of known cell density is shown in Figure D.l. The cell density present in each piece of matrix sample was then determined from the protein concentration compared with the standard curve. n . STANDARD CURVE STANDARD . n

Figure D .l Standard curve of cell density determination by protein assay protein by determination density cell of curve Standard .l D Figure Cell Density (cell/ml) 0 2.00e+6 1.00e+6 . 000+0 0 = .21+ + .19+x * = 0.994 = R*2 2.8139e+4x + 4.1231e+4 = y rti Cnetain (ug/mL) Concentration Protein 20 0 4 0 6 0 8