production of poly(3-hydroxybutyrate) in an airlift bioreactor by ralstonia eutropha

Biochemical Engineering Journal 18 (2004) 21–31

Production of poly(3-hydroxybutyrate) in an airliftbioreactor byRalstonia eutropha

Lana Zanetti Tavares, Elda Sabino da Silva, José Geraldo da Cruz Pradella∗São Paulo Institute for Technological Research, Biotechnology Group, Division of Chemistry,

Avenida Professor Almeida Prado 532, CEP 05508-901 São Paulo, SP, Brazil

Received 11 April 2002; received in revised form 11 April 2003; accepted 16 April 2003

Abstract

The influence of different aeration rates (Qair), 12, 20, 30, 35, 40, and 50 l min−1, on the production of poly(3-hydroxybutyric acid) (PHB)by Ralstonia eutropha DSM 545 during the accumulation phase, was investigated in an airlift bioreactor and the results were compared tothe ones obtained in a stirred tank bioreactor. The time constants for mixing, oxygen transfer, oxygen consumption, power consumption,and kinetic parameters were the tools used to compare the systems. The results showed that, for a superficial gas velocity (Vs) greaterthan or equal to 10 m s−1, the PHB productivity reached 0.6 g l−1 h−1 at 50% of PHB cell content, although, for the velocity of 0.10 m s−1

the observed value of dissolved oxygen concentration in fermentation medium was zero. The analysis of time constants calculated in theaccumulation phase showed that, forVs under 0.11 m s−1, the rate of oxygen consumption was larger than the rate of oxygen transferindicating that this was the rate-limiting step. In the conventional stirred tank bioreactor, the PHB productivity achieved 0.82 g l−1 h−1 at50% of PHB. However, the latter demand higher power consumption than the airlift bioreactor, indicating that, for these bacteria, the lowsupply of oxygen in airlift leads to better performances on PHB production with the advantage of lower demand of energy.© 2003 Elsevier Science B.V. All rights reserved.

Keywords: Airlift bioreactor; Poly(3-hydroxybutyrate);Ralstonia eutropha; Time constants

1. Introduction

Poly(3-hydroxybutyric acid) (PHB) is an intracellular car-bon and energy storage material accumulated by many mi-croorganisms under unfavorable growth conditions such aslimitation of N, P, S, Mg, or O2 and excess of carbon source[1,2]. PHB is a biodegradable thermoplastic polyester thatcan be applied similarly to many conventional petrochemi-cal derived plastics currently in use[3,4].

Ralstonia eutropha has been the most widely used mi-croorganism for the production of PHB because it is easy togrow, it accumulates large amounts of PHB (up to 80% ofdry cell weight) in a simple culture medium and its physi-ology and biochemistry leading to PHB synthesis are wellunderstood[5,6]. The synthesis of PHB usually occurs be-cause the key enzyme for PHB production (�-ketothiolase,Fig. 1) starts to act due to the starvation of nutrients in themedium leading to the accumulation of free Co-A[7]. Ac-cording to the metabolic pathway of the bacteria (Fig. 1),a low concentration of O2 in the medium also leads to an

∗ Corresponding author. Tel.:+55-11-37674315;fax: +55-11-36674055.E-mail address: [email protected] (J.G. da Cruz Pradella).

excess of reduced coenzymes (NADH and NADPH) and ahigher carbon flux could be directed towards PHB synthe-sis for reoxidation of these coenzymes[8] (see step 7 inFig. 1). However, a very severe limitation of oxygen causesformation of intermediates of the Krebs Cycle and even ofthe PHB biosynthetic pathway, harming or even making un-feasible the formation of PHB[9]. It is clear that oxygenlimitation could enhance PHB biosynthesis in recombinantE. coli by decreasing cell growth rate[10]. This also showsthat the timing of PHB biosynthesis can be artificially con-trolled in recombinantE. coli, as the timing of nitrogen limi-tation controls PHB biosynthesis inRalstonia eutropha [11].In other bacteria PHB producers, the same effect of oxygenlimitation could be shown. According to Ward et al.[12]PHB accumulation is induced during oxygen limitation inAzotobacter beijerinckii cultures.

Airlift bioreactors are a special class of pneumatic con-tactors that do not have any mechanical components, likeimpellers and seals. Their main characteristics are low shearstress, simplicity of design and construction[13] and lowenergy requirements for transport rates, besides a better def-inition of internal flow[14,15], and a good aseptic control[16]. The performance of airlift reactors becomes limitedwhen the cell used in the process consumes high amounts

1369-703X/$ – see front matter © 2003 Elsevier Science B.V. All rights reserved.doi:10.1016/S1369-703X(03)00117-7

22 L.Z. Tavares et al. / Biochemical Engineering Journal 18 (2004) 21–31

Nomenclature

C∗ equilibrium concentration of oxygen inthe culture medium (kg m−3)

CER CO2 evolution rate (kg m−3 s−1)CERm mass of produced CO2 (g)kLa volumetric oxygen transfer coefficient (s−1)OUR oxygen uptake rate (kg m−3 s−1)OURm mass of consumed oxygen (g)P ungassed power consumption (Kg m s−1)Pg gassed power consumption (Kg m s−1)PO2 dissolved oxygen concentration (%)Pp productivity of PHB atP/X = 50% of cell

mass (g l−1 h−1)PHBf concentration of PHB at the end of the

fermentation (g l−1)PHB mass of produced PHB (g)PHB% percentage of PHB related to total

biomass (%)S1cons mass of consumed glucose (g)S2cons mass of consumed fructose (g)tc circulation time (s)tmx mixing time (s)trx oxygen consumption time (s)tmt oxygen transfer time (s)Vs superficial gas velocity (m s−1)Xa mass of the residual biomass (g)Yp/s PHB yield from hexoses (g g−1)µXa,max maximum specific residual biomass

growth rate (h−1)µP specific PHB production rate (h−1)

of oxygen. In such cases, special changes on the configura-tion of the sparger and even high compression of air wouldbe necessary[17].

In concentric draught-tube airlift reactors, air is intro-duced into the culture from a sparger at the bottom of thecentral tube. This causes a reduction in the bulk density ofthe liquid in the draught tube (riser) compared to the outerzone of the vessel (downcomer), which sets the culture intocirculation. Continuous production of beer, vinegar, citricacid, and biomass from yeast, bacteria, and fungi has beencarried out in airlift vessels at different working capacities[13]. These bioreactors have also been widely used foranimal cell cultures, especially for antibody production,reaching high productivity[16]. Some authors[18] in theirreview stressed that airlift bioreactors combine high loadingof solid particles and good mass transfer, which are inherentfor three-phase fluidized beds.

Comparison of time constants for different reactor func-tions can be used to determine, which processes are ratelimiting. A small value of a time constant represents a fastprocess[19]. The time constants used in this comparison arethose for oxygen consumption (trx), oxygen transfer (tmt),

and mixing (tmx). The time constantstrx andtmt can be ex-pressed as

trx = C∗

OUR(1)

whereC∗ is the equilibrium concentration of oxygen in theculture medium and OUR, the oxygen uptake rate.

tmt = 1

kLa(2)

wherekLa is the volumetric oxygen transfer coefficient, s−1.Mixing time (tmx) is defined as the time required to reach

some arbitrary level of uniformity in the liquid being mixed.For internal-loop airlift reactors, where gas is sparged intothe riser section, a satisfactory relationship between circu-lation time (tc) and mixing time is[20]

tmx = 3.5tc

(Ad

Ar

)0.5

(3)

whereAd is the downcomer cross-sectional area andAr theriser cross-sectional area.

The circulation time (tc) for internal-loop airlift reactors,considering the relationshipAd/Ar = 0.13–0.56 in waterwith 0.15 M NaCl (surface tension at about 75 mN/m), isdefined as[20]

tc = 4.9

(Ad

Ar

)−0.5

V1/3s (4)

whereVs is the superficial gas velocity defined as the relationof aeration rate,Qair andAr.

The addition ofEq. (4) into Eq. (3)resulted in

tmx = 17.15

(Ad

Ar

)V

1/3s (5)

In this work, we used the time constants as tools to evaluatethe influence of the oxygen supply on the PHB productionby of Ralstonia eutropha DSM 545, during the accumulationphase, in an airlift bioreactor and to assess the suitability ofdifferent superficial gas velocities applied to the riser sectionfor better PHB production performances.

2. Materials and methods

2.1. Bioreactor

All the experiments were performed in a concentricdraught-tube airlift bioreactor (model E10A, B. BraunBiotech International) with a working capacity of 10 l andautomatic temperature and pH controls. The gas spargerwas a stainless steel perforated ring located at the bottomof the riser section. The bioreactor is illustrated inFig. 2.For comparison, a stirred tank fermentor (model ED, B.Braun Biotech International) experiment was performed inthe same conditions. The stirred tank fermentor had 7 l of

L.Z. Tavares et al. / Biochemical Engineering Journal 18 (2004) 21–31 23

Fig. 1. Metabolic pathway of the bacteriaRalstonia eutropha: (1) isocitrate dehydrogenase; (2) �-ketoglutarate dehydrogenase; (3) succinate dehydrogenase;(4) citrate synthase; (5) pyruvate dehydrogenase; (6) �-ketothiolase; (7) acetoacetyl-CoA reductase; (8) PHA synthase. The arrow indicates the pathlimitation.

working volume, an inner diameter of 188 mm and wasequipped with four baffles, three standard flat blade turbineswith a diameter of 75 mm and automatic temperature andpH controls. The gas sparger was a stainless steel perforatedring located at the bottom of the tank.

2.2. Microorganism and media

The strain used wasRalstonia eutropha DSM 545 andthe lyophilized stock culture was maintained at−80◦C. Theseed culture medium composition, the initial medium com-position for batch culture and the feed solution for the accu-

mulation phase are given inTable 1. The trace element so-lution consisted of g l−1 of H2BO4, 0.3; CoCl2·6H2O, 0.2;ZnSO4·7H2O, 0.1; NaMoO4·2H2O, 0.03; NiCl2·6H2O, 0.2;CuSO4·5H2O, 0.01; MnCl2·6H2O, 0.03.

2.3. Microbial PHB production

The cells were cultivated in a two-stage method. First, asuspension of the stock culture in water was inoculated inErlenmeyer flasks (0.1 ml l−1) with nutrient broth medium(Table 1) and incubated for 15 h at 30◦C and 250 rpm. Then,an amount of 10% (v/v) of this culture was inoculated in the


Fig. 2. Airlift bioreactor used in each run.DT is the diameter of the tank(riser and downcomer sections), 0.12 m;HT the height of the tank, 1.00 m;HL the height of the liquid, 0.70 m;DR the diameter of the riser, 0.08 m;sparger type is perforated ring; number of sparger hole, 16; diameter ofsparger hole, 0.8 mm.

seed culture medium and incubated for 15 h at 30◦C and250 rpm.

A volume of about 1.5 l of this culture, large enough toguarantee an initial biomass concentration of 0.4 g l−1, wastransferred to the airlift bioreactor containing about 8 l of ini-tial medium (Table 1). Into the stirred vessel, 1 l of the seedculture and about 6 l of initial medium were transferred. Thetemperature was maintained at 32◦C and the pH was con-trolled at 7.0 with the addition of H2SO4 (4N) and NH4OH(v/v 1:5) solutions. Dissolved oxygen concentration (PO2)was maintained at 20% of air saturation by increasingQairin the airlift reactor and by increasing the agitation speed inthe stirred vessel. An abruptPO2 increase indicate the de-pletion of glucose and fructose from the culture medium atthe end of the growth stage. At that time, PHB accumula-tion was induced by ammonium limitation, by the replace-

Table 1Composition of the medium used in all phases of process

Component Nutrientbroth

Seedculture

Initialmedium

Feedsolution

Glucose (g l−1) – 15 15 190Fructose (g l−1) – 15 15 190KH2PO4 (g l−1) – 1.5 1.29 –Na2HPO4 (g l−1) – 4.45 – –(NH4)2SO4 (g l−1) – 3.0 1.83 –MgSO4·7H2O (g l−1) – 0.2 0.55 –CaCl2·2H2O (g l−1) – 0.01 0.02 –Ammoniac citric ferric

(g l−1)– 0.06 0.05 –

Trace element solution(ml l−1)

– 1 2 –

Meat extract (g l−1) 3.0 – – –Peptone (g l−1) 5.0 – – –

Table 2Conditions of aeration rate in the accumulation phase for each run

Run Qair (l min−1) Vs (m s−1)

1 12 0.042 20 0.073 30 0.104 35 0.125 40 0.136 50 0.17

ment of the NH4OH pH control flask with the NaOH 4N pHcontrol flask and by the addition of a glucose and fructosesolution (190 g l−1) into the fermentor at a constant rate.

PHB production in the airlift was examined by carryingout six different runs under six different values ofQair duringthe accumulation phase according toTable 2.

2.4. Analytical procedures

Cell concentration was determined by measuring the drycell weight after centrifugation at 9220×g for 10 min at4◦C and subsequent filtration using 0.45�m membrane anddrying at 104◦C until constant weight. Residual biomass(Xa) was defined as total cell mass minus PHB mass. ThePHB content (PHB%) was defined as the ratio of PHBconcentration to total cell mass.PO2 was measured with apolarographic DO probe (Ingold, Switzerland). The probewas calibrated after the sterilization process. The 0% oxy-gen was achieved by sparging pure nitrogen into the culturemedium to displace all oxygen in solution and the 100%oxygen was achieved before inoculation when the broth wasfully aerated atQair = 20 l min−1. The PHB concentrationwas determined by gas chromatography[21]. Measurementsof glucose and fructose concentrations in the medium werecarried out with HPLC[21]. The outlet gas compositionwas analyzed by a paramagnetic and an infrared analyzers(Magnos 6G and Uras 10P models of Hartmann Braun).The inlet airflow rate was measured by a mass flowmeter(BRA-002F, Bronkhorst High-Tech).

2.5. Kinetic parameter determination

Although the volume variation throughout the process wassmall, it was taken into account in the kinetic parameterdetermination.

The PHB yield from substrate and PHB productivity, inthe accumulation phase, were defined as follows:

YP/S = �P

�S(6)

PP = �P

tf − ti(7)

where�P is the difference between PHB concentration for-mation at the end and beginning of the fermentation (g).


The specific growth rate was defined as follows:

µXa = 1

Xa

d Xa

dt(8)

The specific PHB production rate was defined as follows:

µp = 1

Xa

dP

dt(9)

where Xa is the residual biomass (g).The OUR, the CO2 evolution rate (CER) and the oxygen

volumetric mass transfer coefficient (kLa) were estimated bythe steady state gas balance around the bioreactor. The inletand outlet gas compositions together with the inlet mass flowrate were calculated as follows:

OUR = xO2 eQair − xO2 sQg

V(10)

wherexO2 e is the inlet mole fraction of O2, xO2 s the outletmole fraction of O2, Qg the outlet gas flow rate (l min−1).

Assuming that the mole fraction of nitrogen in the inletgas is 0.79:

Qg = Qair0.79

1 − xO2 s − xCO2 s(11)

CER= xCO2 sQg − xCO2 eQair

V(12)

wherexCO2 e is the inlet mole fraction of CO2, xCO2 s theoutlet mole fraction of CO2

kLa = OUR

C∗ − CL(13)

where CL is the DO concentration measured by the DOprobe andC∗ the equilibrium DO concentration in theculture medium estimated according to Schumpe, 1985[22].

2.6. Time constants

Equations (1), (2), and (5)were used for the calculationof the time constants. For their determination, for each run,average values of all the variables were used, once there wasa little variation among all of them for each run.

2.7. Power consumption

For the stirred bioreactor the determination of powerconsumption was based on the concepts of the modifiedReynolds number[23]. First, the modified Reynolds number(NRe) was calculated as follows:

NRe = ND2i ρ

µ(14)

whereN is the impeller speed (s−1); Di the flat blade di-ameter, 0.074 m;µ the viscosity of the medium, 0.00135kg m−1 s−1; ρ the density of the medium, 1000 kg m−3. Thevalue ofNRe is 2.4 × 104 for N = 5.83 s−1.

By using the graphic relationship betweenNRe and thepower number (NP), for a six flat blade turbine, the ungassedpower consumed (P) could be determined:

NP = Pg

N3D5i ρ

= 6 (15)

The gassed power consumed in the system was estimated us-ing the aeration number (Na) concept, developed by Ohyamaand Endoh[24], as follows:

Na = Qair

ND3i

(16)

whereQair was the air flow rate, which was maintained at6.67× 10−5 m3 s−1 for the stirred tank experiment.

Then, the ratioPg/P was obtained using the specificgraphic relationship for a flat blade turbine and theNa valuecalculated byEq. (16), leading to the value of gassed powerconsumptionPg, for the stirred tank.

For the airlift bioreactor, the following equation[17] wasused for the power consumption calculation:

Pg = ρLVLVs

1 + (Ad/Ar)(17)

wherePg is the power consumption (kg m s−1), ρL the den-sity of the culture medium, which is considered to be thesame as the water density, 1000 kg/m3. VL the liquid volume(m3), Vs the superficial gas velocity (m s−1).

3. Results and discussion

The growth phase for all experiments lasted approxi-mately 10 h, withPO2 varying from 100% to almost 10%of air saturation (Fig. 9). As can be seen inFigs. 3–8, thegrowth phases were very similar to each other in terms ofsugar consumption and biomass formation. At that time, theintracellular PHB content was between 10 and 25% of cellmass (Fig. 10) and the growth was essentially exponentialat a value ofµXa,max of 0.32 h−1 (data not shown).

During the accumulation phase,Qair influenced theprocess significantly, mainly the PHB biosynthesis perfor-mance, although cell growth was not observed, as Xa wasconstant during this phase (Figs. 3–8).

For the 12 l min−1 experiment (Fig. 3), the mass of PHBincreased from almost 0 up to 32.5 g until the end of growthphase. After that, the amount of PHB continued to increasevery slowly up to a final value of 62.2 g of PHB, or a finalPHB concentration, PHBf = 6.2 g l−1 at 30 h, resulting ina very poor PHB content of 37% of cell mass (Fig. 10).ThePO2 value remained always 0% related to air saturationthroughout the fermentation (Fig. 9), and there was partialconsumption of the fed hexoses (fructose+hexose). Almost189 g of glucose and 152 g of fructose remained at the endof the experiment; having been consumed 282 and 252 g,respectively. About 3.0 g of acetic acid was observed at the


Fig. 3. Kinetic curves for the experiment with aeration rate of 12 l min−1 in accumulation phase: (×) Xa, (�) S1cons, (�) S2 cons, (�) PHB, (�)OURm, and (�) CERm.

end of the run (Fig. 11), indicating severe O2 limitation asobserved by others[9].

The same behavior was observed for the 20 l min−1 ex-periment (Fig. 11). Besides, the PHB mass achieved 88.1 gat the end of run (PHBf = 8.8 g l−1), resulting in a low PHBcontent of 44% of cell mass (Fig. 10). Only a partial con-sumption of the fed hexoses (fructose+ hexose) took place,having been consumed 240 and 306 g of glucose and fruc-tose, respectively.

The PHB biosynthesis derepression appeared to take placein the 30 l min−1 experiment. Up to 256 g of PHB accu-mulated at the end of the run (PHBf = 23.8 g l−1), equiv-alent to a PHB content of 64.3% of cell mass (Fig. 10),

Fig. 4. Kinetic curves for the experiment with aeration rate of 20 l min−1 in accumulation phase, symbols asFig. 3.

despite the fact thatPO2 value remained 0% throughout theaccumulation phase (Fig. 9). Virtually all the added hex-oses, 432 and 440 g of glucose and fructose, respectively,were consumed, and only less than 0.5 g l−1 of glucose re-mained at the end of the experiment. On the other hand,no detectable amount of produced organic acids occurred at30 l min−1.

The profiles of the experiments at 35, 40, and 50 l min−1

were very similar. The PHB cell content attained 71, 74,and 68% of cell mass, respectively (Fig. 10). All the addedhexoses were consumed, no detectable amount of organicacids were observed (Fig. 11) and thePO2 values were notlimiting (Fig. 9).




Kinetics comparison of the experiments showed that, atQair = 12 and 20 l min−1, µp and Yp/s values were lessthan 0.03 h−1 and 0.16 g g−1, respectively (Fig. 12, Table 3).At and above 30 l min−1, an adequate oxygen supply forPHB biosynthesis was reached and it somehow stimulatedthe production of PHB. In these cases,µp attained valuesbetween 0.12 and 0.14 h−1 andYp/s went up to 0.36 g g−1

(Fig. 12, Table 3).In order to explain what really happened in the system, a

relationship among mixing, O2 consumption and O2 transfertime constants, and superficial velocityVs, was established(Fig. 13). According to this figure, when the superficial gasvelocity in the riser section was almost 0.11 m s−1 (at the

Table 3Values ofYp/s andPPHB according to the aeration rate in the accumulationphasea

Qair (l min−1) Yp/s (g g−1) Pp (g l−1 h−1)

12 0.14 –20 0.16 –30 0.29 0.6135 0.30 0.6040 0.32 0.6050 0.36 0.63Stirred tank 0.40 0.82

a Yp/s is the PHB yelled from hexoses,Pp the productivity in PHBcalculated at PHB% equal to 50% of cell mass.




Fig. 9.PO2 values against time of fermentation (t) for each run in the accumulation phase: (�) 12 l min−1; (�) 20 l min−1; (�) 30 l min−1; (×) 35 l min−1;(�) 40 l min−1; (�) 50 l min−1.


Fig. 10. PHB production for each run symbols asFig. 9.

Fig. 11. Evaluation of organic acids in PHB production.

Fig. 12. Specific production rate for each run.


Fig. 13. Time constants as function of superficial velocity.

intersection of the O2 transfer and the corresponding con-sumption time), a state of equilibrium took place betweenthe supply and consumption of oxygen by the bacteria witha time constant around 30 s. Under 0.11 m s−1, the valuesof oxygen transfer time constant always remained below theoxygen consumption time constant values, indicating thatthe system was oxygen limited. Above 0.11 m s−1, the sys-tem was not oxygen limited and an adequate oxygen supplycould be observed.

The results corroborate the fact that, atVs = 0.10 m s−1

(Qair = 30 l min−1), the system was oxygen limited (PO2 =0%), but it did not prevent the PHB accumulation due to themild limited oxygen transfer situation.

Furthermore, this superficial gas velocity influenced thePHB productivity and the yield was almost at the same levelas higher values ofQair (Table 3). It can be explained bythe microorganism metabolic pathway (seeFig. 1). As ob-served by others[9,10,12], a mild O2 limitation at 0.10 m s−1

might have led to an excess of NADH and NADPH reducedcoenzymes and the reoxidation of those coenzymes wouldhave been performed by the formation of PHB. However, avery severe O2 limitation, observed atVs less than or equalto 0.07 m s−1, prevented the formation of PHB, because theenzymes of the metabolic pathways were inhibited by theNADH excess[9], leading to organic acid excretion as shownin Fig. 11.

Although PHB productivity was higher in the stirred tankfermentor (Table 3), the power consumption values (Fig. 14)shows that, in the airlift bioreactor, lower energy was nec-essary for PHB production and even less energy was de-manded at 30 l min−1. So, for economical reasons, the airliftbioreactor might be a better alternative to PHB productionthan the conventional stirred tank bioreactor.

In the present case, the airlift bioreactor may be consid-ered as a completely mixed reactor, because the mixing timeconstants values obtained always remain significantly below

Fig. 14. Power consumption per units of PHB formed vs. time in accu-mulation phase in airlift bioreactor (30 and 50 l min−1) and stirred tank.

the oxygen transfer and oxygen consumption time constantvalues as shown inFig. 13.

4. Conclusion

The effect of the superficial gas velocity on the PHBproduction has been investigated in an airlift bioreactor.The PHB productivity and PHB yield from carbon sourceincreased with superficial gas velocity in the accumulationphase. However, the superficial gas velocity 0.10 m s−1

(Qair = 30 l min−1) was adequate enough to guarantee al-most the same PHB production yields as higherVs values.The time constant analysis corroborate the fact that theoxygen consumption rate was the process-limiting step atsuperficial gas velocities higher than 0.11 ms−1. Thus, theadvantage of the airlift bioreactor over the mechanical stirredtank may be the lower energy demand for PHB production.Experiments using fed-batch strategies have been performed


to achieve high cell concentrations and PHB contents inairlift bioreactors in order to decrease production costs.

Acknowledgements

The author Lana Zanetti Tavares would like to thankCapes and FAPESP for the financial support. This work wassupported by FINEP/MCT.

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production of poly(3-hydroxybutyrate) in an airlift bioreactor by ralstonia eutropha

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