Biotechno ELSEVIER
Journal
of Biotechnology
52 (1996) 21-39
Experimental identification of a scalable reactor configuration for lignin peroxidase production by Phanerochaete chrysosporiuwl Francesca Diprtimwto
di Scirnx
dci Materiali
Bosco, Bernard0 c’Iqegnerio
Received 20 May 1996; revised
Ruggeri”,
Chirnic,u Politrrwco Torino. Irtr!,,
14 September
Guido
Sassi
di Torino. C. so Duw
1996; accepted
I5 September
dqli
Ahruz~i 24 l012Y.
1996
Abstract The present work concerns the experimental verification of a suitable scalable reactor configuration for lignin ~/z~~~,so.spoviunz. Tests were performed in 30 ml working volume Erlenmayer peroxidases production by Phanerochaetr flasks, 1.50 ml Fernbach flasks and a 5 cm ID, 83 cm high Trickle Fixed Bed Reactor working in batch mode with 250 ml of liquid medium. Shear stress, biofilm thickness and aeration have been shown to be the key parameters for the control of lignin peroxidase production. Biofilm thickness was identified as a main scale up parameter. The results, compared with those found in the literature, are encouraging for further optimization of the reactor device and scale increase. Biofilm thickness and gas and liquid superficial velocities have been identified as variables in the optimization of lignin peroxidase production in a Trikle Fixed Bed Reactor configuration. Copyright ‘0 1996 Elsevier Science B.V. Kqwords:
Lignin peroxidase;
Phanrrochaetr
chryso.sporium;
1. Introduction
Phanerocharie chrym~porium is a well known ligninolytic white-rot fungus that is able to degrade and, in some cases, mineralize several xenobiotic compounds including polychlorinated biphenyl (Eaton, 1985; Morris and Lester, 1994), ___
* Corresponding author. 016%1656/96/$15.00
E-mail:
[email protected]
Copyright
PII SO 168- 1656(96)01620-3
Scale-up;
Immobilized
biomass;
Fixed bed reactor
DDT (Bumpus et al., 1985; Bumpus and Aust. 1987), dioxin (Hammel et al., 1986), benzo(a)pyrene (Haemmerli et al., 1986) and trinitrotoluene (Fernando et al., 1990). An extracellular ligninolytic system is involved in the initial steps of degradation of many of these compounds. Many authors have highlighted the high sensitivity of P. chrysosporium to agitation and aeration: the highest lignin peroxidase (LIP) productions have been obtained in surface culture
cl 1996 Elsevier Science B.V. All rights
reserved
(thin mycelial layer) using unstirred very small reactors (Kirk et al., 1978; Faison and Kirk, 1985; Leisola et al., 1987; Venkatadri and Ivine, 1990). This may be due to the low stresses in such reactors coupled with good oxygen availability (directly from the gas phase) to a thin biofilm. The effect of aeration was elucidated in the works of Bar-Lev and Kirk (1981) and Faison and Kirk (1985) and according to these authors, both lignin mineralization and LIP synthesis are increased in cultures grown under high O2 tension. The effect of a solid support on which biomass may grow has been considered to be positive for exoenzyme synthesis (Bonnarme et al., 1991). The scale up of such a reactor configuration is a hard task because of the poor controllability of such a system, its geometrical attributes and the strong interaction between the biotic and abiotic phases. Many efforts have been devoted to achieve a scalable reactor configuration giving acceptable performance (Michel et al.. 1990; Janshekar and Fiechter, 1988; Linko, 1988; Willershausen et al., 1987; Kirk et al., 1986; Leisola and Fiechter, 1985). The negative effect of agitation. due to the shear stress effects, on the production and secretion of ligninolytic enzymes was reported by Michel et al. (1990), who obtained a maximum value of LIP production of about 0.175 U ml - ’in a stirred tank reactor configuration. Bonnarme et al. (1993) obtained a higher LiP production in a pneumatically agitated bioreactor than in mechanically agitated ones. In the former case the low shear-stress provided a very efficient mycelium. In submerged bioreactors, biomass grows as pellet colonies completely immersed in the culture medium without direct contact with the gas phase and this may induce oxygen mass transfer limitation both inside and outside the pellets (Wittler et al.. 1986; Kobayashi et al., 1973). These considerations suggest the necessity of designing a suitable bioreactor configuration that provides good oxygen transfer in a low-shear environment. In order to identify scale up criteria and the parameters that affect LIP production, the behaviour of 30 ml and 150 ml unstirred surface culture bioreactors is discussed here. Since they are suitable for scale-up purposes, in the present
study a Trickle Fixed Bed Reactor (TFBR) has been tested in order to control the thickness of biofilm and contact area between mycelium and both gas and culture medium and to avoid diffusive mass transfer limitation. To achieve a suitable spatial distribution of biomass on the bed, the inoculum was entrapped in Ca-alginate adhering to an inert support (ceramic Berl saddles) used as a random packing in the TFBR.
2. Materials
and methods
2.1. Strain and culture conditions Phanerocharte chr~~soqmrium BKM-F- 1767 was maintained on 2% malt extract agar plates. The cultures were inoculated as 5 x 10”’ conidia l- ’ suspension in distilled water directly added to the medium or encapsulated in alginate to give a concentration of 5 x 10” 1~ ’ conidia referred to the liquid medium. The cultures were grown at 39°C. A nitrogen limited medium with glucose as the carbon source was used, prepared according to Tien and Kirk (1988). The initial pH value was adjusted to 5.5. Asolectin from soyabean (11145 Fluka) was added (0.75 g 1~ ‘), but no veratrylalcohol (3,4_dimethoxybenzyl alcohol) was introduced in the initial medium. After about 48 h of reaction, the temperature was changed to 30°C and veratrylalcohol (2.5 mM) and asolectin from soyabean (0.1 g 1~ ‘) were added to the liquid medium. Pure oxygen was flushed through the reactor to maintain high concentrations in both gas and liquid phases. This procedure has been previously assessed (Bosco, 1995).
2.2. Remtor conditions
configurations
md operational
Tests were carried out in three different reactor configurations: 500 ml Erlenmeyer flasks, 2800 ml Fernbach flasks, and a 5 cm ID fixed bed reactor (TFBR). Two kinds of support were used: ceramic Raschig rings 6.35 mm (114”) and ceramic Berl saddles 12.7 mm (l/2”).
2.2.1. Erlenmeyer flasks The first configuration was 500 ml volume (75 cm’ bottom area) with 41 g of Raschig rings as a single layer, 30 ml of culture medium and oxygen flushed daily at 600 ml min ’ for 2 min. The system was not stirred. The test conditions were deduced from Bosco (1995).
2.2.2. Fernbach jiasks The second configuration was 2800 ml volume (375 cm2 bottom area) with 192 g of Raschig rings as a single layer, 150 ml of culture medium and oxygen flushed daily at 600 ml min -’ for 12 min. This system was also not stirred.
2.2.3. TFBR A glass vessel 5 cm ID, 83 cm high, the bottom being used as a reservoir, with 300 g of ceramic support, 250 ml of culture medium circulated (down flow) batchwise around an external loop using a peristaltic pump at 30 ml min _ ‘. Fig. 1 shows a schematic drawing of the reactor. In this TFBR configuration the following runs were carried out:
1
Fixed Bed Reactor
2.3. Analytical
methods
23.1. Analytical assays Glucose and ammonium concentrations in the extracellular fluid were determined spectrophotoenzymatic kits (Boehringermetrically, using Mannheim No. 716251 and No. 1112732, respectively). 2.3.2. Enzymatic assays Lignin peroxidase activity was determined spectrophotometrically in the extracellular fluid undiluted, by the method of Tien and Kirk (1988) at 310 nm and 39°C with veratrylalcohol as a substrate. Mn(II)-dependent peroxidase activity was determined spectrophotometrically at 22°C and 568 nm, by the method of Paszczynski et al. (1988) with 2,6-dimethoxy-phenol as the substrate.
1. Fixed bed 2. Gas compressor 3. Liquid pump 45. Gas filters 6. Gas humidifier 7. Liquid distribution 8. Sample port 9. Liquid filter
Fig. I. Trickle
Run 1: Raschig rings were used and oxygen was flushed daily at 600 ml min _ ’ for 15 min. Run 2: Raschig rings were used and oxygen was flushed continuously at 1S-60 ml min _ ’. Run 3: Raschig rings were used and the medium was circulated at 30 ml min ~_‘. Oxygen was flushed continuously at 5-15 ml min ’ and the gas phase was partially recirculated to obtain a flow rate of 80 ml min _ ’ inside the column. Run 4: Berl saddles were used and oxygen was flushed continuously at 15 ml min with the gas phase partially recirculated to obtain a Aow rate of 80 ml min- ’inside the column. Inoculum was entrapped in Ca-alginate (2%w,/v Na-alginate, 2%/v Tween80, O.S’%v/v Linoleic acid, ionotropitally gelled in a 0.5 M CaCl, solution). Conidia were inoculated in the alginate solution before gellation. The gellation procedure was deduced from previous tests (Ruggeri et al., 1991).
(TFBR)
device.
2.3.3. Biomass determination Immobilized mycelium was measured in terms of dry weight after drying to constant weight at 105°C on a Petri dish. The predetermined weight of support (Raschig rings or Berl saddles) was subtracted from the weight of support plus mycelium
_
12r
The amount of total protein in the extracellular fluid was determined by the method of Bradford (1976, Bio-Rad kit) using bovine serum albumin as a standard.
??
A
x
10
A 0
8
Glucose, g I -’ Ammottittm x100, g DH Biomass, g I -’
I -’
I
:\
.
3. Results and discussion
In Erlenmeyer flasks, 41 g of l/4” Raschig rings corresponded to a single layer of supports which needed 30 ml of medium to be completely immersed. Biomass grew as a biofilm at the gas/liquid interface: 13 f 2 g m ~ ’ of biomass (as dry weight) per available surface area were measured at the end of 12 separate runs (based on the bottom area of the flask as an estimate of the surface area available for biomass growth). A maximum LIP activity of 0.4-0.6 U ml ’ was found. Fig. 2 shows the results obtained in this Erlenmayer flask device (using different flasks sacrificed along the time course of the run). A lag phase without any LIP activity was observed (100 h), and the maximum activity was observed after 100 h from its first appearance. Manganese peroxidase (MnP) activity was observed at between 50 and 200 h but at very low levels (O.OZZO.03 U ml- ‘). The glucose consumption rate was almost constant over the whole run and may be estimated to be 0.025 g 1~ ’ h- ‘. Ammonium is exhausted during the first 30 h while the pH decreases from 5.5 to 4.5 during the first 50 h and then remains almost constant. Biomass determination were subject to high variability mainly due to the experimental methods adopted. The mycelial growth is restricted to the first 80 h; a small reduction then occurs to a constant value which is then maintained over the whole period during which LiP activity persists. The total protein concentration profile vs. time is similar to that of LiP while the substrate is available; when the substrate concentration is below 1 g 1~ ’ the total protein concentration increased and LIP activity decreased. This observation can be due to protease production (Dey et al.. 1991; Eriksson and Petterson, 1988) which can explain also the sharp de-
time, h 0.025
u ”
A I
I
Total protein,. g MnP, U ml .’
TWO
0.5
I -’
I 0 Lip.
U ml -’
500
time, h
Fig. 2. Erlcnmeyer
flasks run
crease of biomass concentration which almost disappears in 24 h. As shown in Fig. 2. the maximum LIP activity occurs when the biomass concentration and glucose consumption rate are almost constant. Bearing in mind that, in the adopted experimental conditions, the bioreaction is at the surface and any shear stress due to agitation is absent, one can argue that LIP production might be affected by the biofilm thickness. Therefore a suitable scale up criterion might be to try to maintain the same thickness of biomass layer. This can be satisfied by keeping constant the ratio between the medium volume and the support area for mycelium growth at the gas/liquid interface. The biomass (as dry weight) per unit of surface area available for its growth may be considered as the parameter which represents the biofilm thickness. Experiments carried out with Fernbach flasks were designed to test this hypothesis; the flasks were filled with a
single layer of support completely covered by the culture medium. These two conditions could create the conditions that match the criterion. Biomass grew as a biofilm at the gas/liquid inter’ of biomass (as dry weight) faceand 15$2gm per available surface area were measured at the end of four separate runs (based on the bottom area of the flask as an estimation of the surface area available for biomass growth). The value of the biomass superficial concentration was close (in the range of experimental errorj to that obtained in the Erlenmeyer flasks. Fig. 3 shows the results obtained in the Fernbach flask devices (using samples withdrawn daily). The behaviour vs. time of the LiP activity was similar to that observed in the Erlenmeyer flasks, with a maximum activity 25% higher (repeated tests gave a maximum activity of 0.550.7 U ml ~ ‘) The glucose consumption rate increased to 0.029 g 1~~’ h ’ and glucose was reintroduced after 200 h in order to avoid protease production. The results show that, keeping constant the ratio between medium volume and available area for mycelium growth, the criterion was substantially matched. The criterion may be considered right and the small increase of performances might be explained by the lower variability observed in the biofilm thickness.
I-
Glucose, g I
-’
10
time, h Fig. 3. Fernbach
flask run
LIP, U ml -’
In order to identify a suitable scalable reactor configuration, it is necessary to take into account constraints already identified in the literature concerning the main phenomena involved (Linko: 198X; Capdevila et al., 1989: Asther et al., 1990; Bonnarme et al.. 1991, 1993) as well as the experimental expertize already acquired in the laboratory tests (Bosco, 1995). The system device must have the following characteristics: supported biofilm; low mechanical shear stress on the biofilm due to liquid and gas fluid dynamics; large support area available for biofilm growth; large liquid/gas/mycelial interfaces. Furthermore, taking into consideration the above results. the biofilm thickness and gas/liquid/biofilm mass transfer rate should be controllable. A first criterion generally followed for chemical reactor scale-up is to maintain geometrical similarity. A plate reactor such as that used for citric acid production (Rohr et al., 1983) might be considered but this device has very great problems in handling and in the control of its performance: liquid and gas phase mass transfer are hardly controllable; the ratio of support to liquid volume is completely determined by the support dimension; a continuous operational mode is not available. These consideration become more stringent as the scale increases and, for these reasons, geometrical similarity is not maintained in this study. By comparison, a device such as a TFBR (Trickle Fixed Bed Reactor) seems to match many of the above constraints: a packing may support the biomass, percolating liquid has a very low shear effect, the surface area for mycelial growth and gas/liquid/mycelial interfaces is large and easily enhanced by using different types of support. Biofilm thickness may be controlled by varying the ratio of liquid volume to available packing surface with a constant medium composition. Gas/liquid/biofilm mass transfer rates may be controlled by varying gas and liquid superficial velocity. Using a TFBR. the main problem that might arise in the control of biomass distribution as a consequence of imperfect liquid and inoculum distribution.
26
F. Bosco et ul.
IT-
: Journal oj’Biotechnology’ 52 (1996) 21-29
In order to scale up from the flask system to a fixed bed device with constant biofilm thickness, the biomass has been presumed to grow on 40% of the external surface of the Raschig rings, as practically it was in the surface reactor devices. This surface projection was used to calculate the medium volume assuming a homogeneous distribution of biomass in the bed volume, in order to achieve the above scale up criterion. To avoid changes in kinetic behaviour due to substrate and micro nutrient limitation and/or inhibition, the medium composition was the same as that used in the flask runs. Initially pure oxygen was flushed daily (as had been done in the surface reactor devices); the liquid phase was circulated at low superficial velocity (0.6 cm min ‘) to avoid shear stress, although bed wetting was limited to lo20% as estimated by Coulson and Richardson’s procedure (Coulson and Richardson, 1978). A 5 cm diameter column was chosen to limit the scale increase with a reasonable bed height and at the same time to reduce liquid distribution problems at the top of the reactor. During Run I biomass grew only at the top and at the bottom of the support bed where reservoirs of oxygen were present: this suggested a need to supply oxygen continuously through the bed to avoid macroscopic oxygen mass transfer limitation. In Run 2 the superficial gas velocity was initially set at 0.9 cm min ’ and then increased up to 3 cm min-’ after 150 h; the results are shown in Fig. 4. The behaviour vs. time of the LIP activity was qualitatively similar to that obtained in the flask device tests, but the measured LIP activity was very much lower with respect to both maximum value and persistence. Furthermore. qualitative observation indicated that biomass was badly distributed in the bed and growing also in the reservoir so that it gave some trouble in the liquid distribution system. In Run 3 oxygen was fed continuously at 5 ml min ’ and recirculated in order to give a superficial gas velocity of 4 cm min -’ inside the reactor. In this way oxygen concentration and superficial gas velocity could be controlled by separate operational parameters. A rough filter was inserted in the liquid circulation system to avoid obstruction of the liquid distribution device and of the top of
Glucose,
g I
-’
??
-’
x Ammonium x 100, &I
12
Lip, U ml -’
-
6
0.06
time, h
Fig. 4. Run 2 in TFBR
with 114” Raschig
rings
the bed caused by free mycelium circulating with the liquid. During the run the oxygen flow rate was increased up to 10 ml min ’ after 110 h and 15 ml min - ’ after 220 h. The results for Run 3 are shown in Fig. 5; LIP activity was enhanced, but it was still very much lower than that obtained in flask devices. The biomass distribution was also found by visual observation to be inadequate and flooding of the bed appeared. coupled with maldistribution of the gas and liquid flows. The glucose consumption rate (0.043 g 1-l h ~ ‘) was greater than in the flask devices; this may be I ??
x
Glucose, g I.’ Ammonium x 100, g I
0
LIP, U ml -’
J
200
300
time, h
Fig. 5. Run 3 in TFBR
with l/4” Raschig
rings
1 1.6 - 1.4
? ?LIP, U
d-l
- 1.2 - 1.0
0.x - 0.6
0 Fig 6. Run encapsulated
100
200
4 in TFBR conidia.
300
with
weight) per available surface area measured. No biomass was observed to grow free in the liquid medium. A glucose consumption rate lower than that in Run 3 was measured (0.031 g 1~ ’ h- ‘), although this was still greater than those measured in the flasks runs. Fig. 7 shows the average LIP activity of repeated Fernbach flask tests compared with that obtained in the TFBR device. The LIP activity is higher in the TFBR over the whole duration of the runs and has a longer persistence at over 0.25 U ml _ ‘.
400 50 Time, h I$!” Bcrl saddles,
Ca-alginate
due to the better liquid mixing reducing the external mass transfer resistance. By this test, the final consideration identified is that the main problem arises because of inadequately distributed mycelium growth on the supports and this may derive mainly from the maldistribution of liquid flow containing the inoculum and the low strength of the link between biomass and support. To avoid flooding, the support was substituted by 112” Berl saddles already tested in a Fernbach flask device (Bosco, 1995), so that the packing channel diameter was doubled. This resolved the flooding troubles, but it increased liquid and hence biomass maldistribution. A better inoculum distribution was achieved by entrapping conidia in a Ca-alginate matrix immobilized on the supports. The whole surface of the saddles was covered by the polymer. Biomass was assumed to grow on 30% of the surface in order to calculate the ratio between the medium volume and the support area for mycelium growth taking into account biomass maldistribution and the surface not available because of contact interaction between saddles inside the bed. Fig. 6 shows that the results of Run 4 quantitatively match that obtained by flask devices with a greater LIP activity and a longer persistence of high LIP activity. Biomass appeared to be well distributed over the bed volume; the 25-40% of saddle surface was found to be covered by biomass with 8- 13 g m _ ’ of biomass (as dry
4. Conclusions Shear stress. biofilm thickness and aeration have been observed to be the key parameters for the control of LIP production by Phunerochaete chrysosporium. The ratio between surface area available for biomass growth and culture medium volume, i.e. biofilm thickness, has been verified to be a main scale up parameter. A Trickle Fixed Bed Reactor configuration, with conidia distribution on the support surface by a Ca-alginate entrapment procedure has been shown to be scalable using well consolidated procedures of chemical reaction engineering. It shows a good control of the biofilm thickness coupled with a low shear stress on the biomass. It guarantees gasjliquidjbiophase contact similar to that observed in simple surface cultures and good control of the transport phenomena. The growth on 1.J
Lip. U ml
I
*
,
I
I
-I 1.2
r
Time, h Fig. 7. LiP production in Fernbach flasks and in TFBR Berl saddles and Ca-alginate encapsulated conidia.
with
28 Table I LiP production
in bioreactor
Bioreactor configuration
Volume
(1)
Maximum LiP productivity (U I ’hh’ )
Reference
0.017 0.130 0.232 0.102
0. I08
Maximum (U ml-‘)
LiP activity
Air Lift Rotating contactor Stirred tank Stirred tank
F,’ Ih F F
Bubble column Stirred tank Bubble column
F F
7.5 0.50 5.0
0.730 0.175 0.760
5.40 I .08 5.43
Leisola and Fiechter, 1985 Kirk et al., 1986 Willcrshausen et al., 1987 Janshekar and Fiechter. 1988 Linko. I988 Michel et al.. 1990 Bonnarme and Jeffries,
Fixed bed
I
0.25
I.220
6.26
This work
I
I.5 2.5 1.0 42
I .08 1.44 0.32
1990
,‘Free biomass. “Immobilized biomass
solid supports provides a mechanical resistance to the biofilm structure and furthermore, alginate enhances the biomass to support bonding. The LIP activity obtained in this TFBR configuration was greater than that reached in surface cultures in flasks. The enhancement of glucose consumption rate from flask to TFBR device suggests the possibility of mass transfer limitation in the liquid phase during runs in the non agitated flasks. The liquid mass transfer resistance seems to limit glucose consumption rate but it is not demonstrated as affecting the level of LIP production; nevertheless this could explain the greater LiP activity seen in the TFBR configuration in comparison to the flask devices. On the other hand the biofilm thickness has been estimated to be lower in the TFBR than in the flask reactors and this too could explain the enhancement of LIP production. As shown in Table 1, LiP activities and LIP productivity reached by the TFBR device have been greater than that obtained by other authors with the same microorganism and similar culture medium (glucose as carbon source and nitrogenlimited). Using another strain and/or different culture medium, higher LiP activities have been obtained (Asther et al.. 1990; Bonnarme et al., 1993). However, the TFBR seems to be more promising because of its better control of the key parameters. The operational parameters (Liquid
and gas flow rates, inoculum concentration and biomass concentration) have not been optimized in order to assure the best LIP productivity. Research in progress is intended to study in a systematic way the influence of dynamic parameters affecting the TFBR performance in order to facilitate a bigger reactor for LIP production suitable for bioremediation applications. Biofilm thickness, gaslliquidibiofilm interface area, gas and liquid superficial velocity will be considered as variables in the optimization of lignin peroxidase production in a Trikle Fixed Bed Reactor configuration. References Asther. M.. Bellon-Fontaine, M.N., Capdevila, C. and Corrieu, G. (1990) A thermodynamic model to predict Phnrzerochcww c,hv~.so,sporilml IKA-I 2 adhcbion to various solid carriers in relation to lignin peroxidase production. Biotechnol. Bioeng. 35. 477~ 482. Bar-Lev. S. and Kirk. T.K. (1981) The effect of molecular oxygen on lignin degradation by P/wwroc~/uctr cl~r~~so.sporium. Biochem. Biophys. Res. Comm. 99, 373 378. Bonnarme. P. and Jeffries, T.W. (1990) Selecti\,e production of cxtracellular peroxidases from P/mwroc/~~err c/qm~~porimz in an Airlift bioreactor. J. Ferment. Bioeng. 70, 158-163. Bonnarme, P.. Delattrc, M.. Corrieu G. and Asther, M. (1991) Peroxidase secretion by pellets or immobilized cells of Phcuwroc~/ucte c,h~~s~~~~~~~~~~~~?~ BKM-F-l 767 and INA- 2 in relation to organellc content. Enzyme Microb. Tcchnol. 13, 7277733.
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Bonnarme. Asther, ganese
of Biotechnology
P., Delattre, M., Drovet, H., Corrieu, G. and M. (1993) Toward a control of lignin and manperoxidases hypersecretion by Phanerochaete chrysosporium in agitated vessels: evidence of the superiority of pneumatic bioreactors on mechanically agitated bioreactors. Biotechnol. Bioeng. 41, 440-450. Bosco, F. (1995) Valutazione sperimentale dei parametri di scale up per la produzione di Lips con Phanerochaete chrysosporium Burds. Ph.D. Thesis. University of Torino, Torino, Italy. Bradford, M.M. (1976) A rapid and sensitive method for the quantitation of microgram quantities of protein utilizing the principle of protein dye binding. Anal. Biochem. 72, 2488254. Bumpus, J.A., Tien, M.. Wright. D. and Aust, S.D. (1985) Oxidation of persistent environmentally pollutants by a white rot fungus. Science 228, 143441436. Bumpus, J.A. and Aust, S.D. (1987) Biodegradation of envronmental pollutants by the vvhite rot fungus Phunerochaete chrysosporium : involvement of the lignin degrading system. Bioassay 6, 166 170. Capdevila, C., Corrieu, G. and Asther. M. (1989) A feed-harvest culturing method to improve lignin peroxidase proPhanerochaete INAduction by chryosporium immobilized on polyurethane foam. J. Ferment. Bioengng. 68. 60-63. Coulson, J.M. and Richardson, J.F. (1978) Chemical Engineering, 3rd Edn., Vol. 2. Pergamon Press, Oxford, pp. 163. Dey, S., Maiti, T.K., Saha, N.. Banerjee, R. and Bhatprotease and amylase tacharyya, B.C. (I 99 I ) Extracellular activities in ligninase producing liquid culture of Phanerochaete chrysosporium. Process Biochem. 26, 325- 329. Eaton, D.C. (1985) Mineralization of polychlorinated chr~sosporium: a ligninolytic biphenyls by Phnnrrochaetr fungus. Enzyme Microb. Technol. 7, 1944196. Eriksson, K.E. and Petterson, B. (1988) Acid proteases from Sporotriclum pulvurulentum. Methods Enzymol. 59, 500.m 508. Faison, B.D. and Kirk, T.K. (1985) Factors involved in the regulation of a ligninase activity in Phunerockaete chrysoqwium. Appl. Environ. Microbial. 49. 2999304. Fernando, T.. Bumpus. J.A. and Aust, S.D. (1990) Biodegradation of TNT (2,4,6-trinitrotoluene) by Phaoerodzarte chrysospporium. Appl. Environ. Microbial. 56, 166661671. Haemmerli, SD., Leisola, M.S.A., Sanglard, D. and Fiechter. A. (1986) Oxidation of benzo(a)pyrene by extracellular ligninase of Phanerorhaete chrysosporium. J. Biol. Chem. 261, 16900&16903. Hammel. K.E., Kalyanaraman, B. and Kirk, T.K. (1986) Oxidation of polycyclic aromatic hydrocarbons and dibenzo(p)-dioxins by Phanerochaere chrywsporuwn ligninase. J. Biol. Chem. 261, 16948816952. Janshekar, H. and Fiechter. A. (1988) Cultivation of Phanerochaete chrysosporium and production of lignin peroxi-
52 (1996)
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