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Industrial Crops and Products 89 (2016) 478–485

Contents lists available at ScienceDirect

Industrial Crops and Products

jo u r n al homep age: www.elsev ier .com/ locate / indcrop

ingle cell oil production integrated to a sugarcane-mill: Conceptualesign, process specifications and economic analysis using molassess raw material

.P.F. Vieiraa,b, J.L. Ienczakb, P.S. Costaa,b, C.E.V. Rossellb, T.T. Francoa, J.G.C. Pradellab,∗

School of Chemical Engineering University of Campinas, UNICAMP, CEP 13083-852 Campinas, São Paulo, BrazilBrazilian Laboratory of Science and Technology of Bioethanol (CTBE), Brazilian Center for Research in Energy and Materials (CNPEM) Rua Giuseppeáximo Scolfaro, 10.000–Polo II de Alta Tecnologia, Caixa Postal 6192, 13083-970 Campinas, SP, Brazil

r t i c l e i n f o

rticle history:eceived 26 January 2016eceived in revised form 21 May 2016ccepted 25 May 2016vailable online 10 June 2016

eywords:ingle cell oilugarcane mill integration

a b s t r a c t

This study aimed to assess the single-cell oil production integrated to a sugarcane-mill and molassesas main raw material. Rhodotorula glutinis lipid production with low-cost raw sugarcane molasses wasaccessed. The envisage process relies in an optimized sugarcane molasses culture media in a fed-batchstrategy to achieve a high-cell density high-lipid cell content. A Process Flow Design (PFD) integratedto a typical Brazilian sugar mill was proposed according to fermentation experiments. Performed eco-nomic analysis provided data for discussion on strategic equipment that can facilitate the viability to theprocess. Typical results of established protocol were total biomass concentration 62.25 g l−1, lipid produc-tivity 0.42 g l−1 h−1, sugar conversion yield to lipids 0.21 g of lipid g−1 of total reducing sugar. Microbial

olassesrocess designconomic analysis

lipid production and defatted biomass plant production was proposed at a nominal capacity of respec-tively 16,720 ton and 21,600 ton per year, running 8300 h per year. The projected selling prices of theseproducts were US$ 1,300.00/ton of lipid and US$ 500.00/ton of defatted yeast. It was demonstrated thatthe industrial plant was potentially attractive when capital investment cost was decreased with the useof low cost epoxy-lined carbon steel stirred bioreactor. In this alternative, the internal rate of return (IRR)will be 24.61% leading the NPV (at 7% interest rate) of US$ 67,797,000/year (before taxes).

. Introduction

The fossil fuels shortage and lack of sustainability raised ques-ion on its usage in a long-term basis. They are responsible toncrease CO2 level in the atmosphere which is directly associated

ith global warming. It is a consensus that global energy towards tonergy from renewable sources and consequently, there is renewednterest in the production and use of fuels from plants or organic

aste (Macrelli et al., 2012).Production of a cheap renewable fuels to replace fossil fuels is

ubject in the political agendas of many countries, aimed at theevelopment of a reliable energy source to ensure fuel security, pro-

ote rural development and to address climate change by reducing

reenhouse gases emission (Macrelli et al., 2012).Among proposed renewable fuel, biodiesel (methyl and/or ethyl

sters of long chain fatty acids of vegetable oils) presents as the fuel

∗ Corresponding author.E-mail address: jpradella@bioetanol.org.br (J.G.C. Pradella).

ttp://dx.doi.org/10.1016/j.indcrop.2016.05.046926-6690/© 2016 Elsevier B.V. All rights reserved.

© 2016 Elsevier B.V. All rights reserved.

with the highest potential to replace diesel oil (Leung et al., 2010).Usually the industry adopted for biodiesel production the transes-terification reactions of vegetable oils (soybean oils, rapeseed oils,palm oils, and its waste cooking oils) with methanol or ethanol.Vegetal oil high price and low vegetable oil productivity has beinghindered their use in large scale as its production costs are associ-ated with raw material mainly with raw-material price up to 80%of its value (Haas et al., 2006). In recent years, much attention hasbeen done to the exploration of microbial oils, which might becomeone potential oil source for biodiesel production (Li et al., 2007). Oilsfrom microorganisms, known as single cell oils (SCO) emerge as apotential substitute for vegetable oils and presents 80–90% of thetriglycerides content (Ratledge, 1991).

On the other hand, SCO has become an important source ofpolyunsaturated fatty acids (PUFA). PUFA are present in cellularmembrane of neurons, sensory cells, sub-cellular organelles and

play important functional role as precursors of prostaglandins,lipoxins and leukotrienes. Moreover, specific PUFAs are rec-ommended for the prevention and treatment of cardiovascularand inflammatory diseases, brain disorders and obesity. Rele-

J.P.F. Vieira et al. / Industrial Crops and Products 89 (2016) 478–485 479

Table 1Kinetic parameters for some high-cell-density Rhodotorula glutinis and Rhodosporodium toruloides (Rhodotorula glutinis anamorph state) cultivation in bioreactor aiming tolipid production.

Yeasts �max (h−1) Substrate Xtotal (g l−1) Lipid (%; w w−1) Pr (g l−1 h−1) Reference

R. glutinis 0.235 Sugarcane molasses 62.2 54 0.42 Improved fed strategy; this studyCCT 2182R. glutinis 0.260 Glucose 185.0 40 0.88 Pan et al. (1986)NRRL Y 1091R. toruloides Y4 0.130 Glucose 106.5 67 0.54 Li et al. (2007)R. toruloides Y4 – Glucose 127.3 62 0.57 Li et al. (2007)R. glutinis – Corncob hydrolysate 70.8 47 0.17 Liu et al. (2015)CGMCC 2.703

Table 2Equipment list of microbial lipid production.

Name Type Units Size Material Purchase Cost ($/Unit)

V-101 Receiver Tank 1 10.00 m3 CS 33,000HX-101 Heat Exchanger 1 70.00 m2 CS 104,000PM-101 Centrifugal Pump 1 0.25 kW SS316 11,000V-102 Flash Drum 1 0.20 m3 CS 3000HX-105 Heat Exchanger 1 50.00 m2 CS 85,000V-103 Receiver Tank 1 5.00 m3 CS 20,000PM-102 Centrifugal Pump 1 0.22 kW SS316 10,000V-104 Receiver Tank 1 10.00 m3 CS 33,000HX-102 Heat Exchanger 1 500.00 m2 CS 339,000PM-103 Centrifugal Pump 1 2.05 kW SS316 27,000V-105 Flash Drum 1 1.00 m3 CS 8000HX-103 Heat Exchanger 1 50.00 m2 CS 85,000V-106 Receiver Tank 1 10.00 m3 CS 33,000PM-104 Centrifugal Pump 1 1.87 kW SS316 26,000R-101 Stirred Reactor 3 2.00 m3 SS316 618,000R-102 Stirred Reactor 3 45.00 m3 SS316 1,054,000R-103 Stirred Reactor 11 500.00 m3 SS316 2,818,000PM-105 Centrifugal Pump 1 0.01 kW SS316 10,000PM-106 Centrifugal Pump 1 0.22 kW SS316 10,000CF-101 Centritech Centrifuge 1 55.00 m3/h SS316 344,000V-109 Receiver Tank 1 5.00 m3 CS 20,000HX-104 Heat Exchanger 1 155.00 m2 CS 168,000PM-108 Centrifugal Pump 1 0.76 kW SS316 17,000V-110 Flash Drum 1 1.00 m3 CS 8000SDR-101 Spray Dryer 1 130.41 m3 SS316 312,000HX-106 Heat Exchanger 1 15.00 m2 CS 41,000M-101 Centrifugal Fan 1 55,292.91 m3/h CS 19,000PM-107 Centrifugal Pump 1 0.70 kW SS316 17,000PM-109 Centrifugal Pump 1 2.05 kW SS316 27,000HG-101 Homogenizer 1 40.00 m3/h SS316 161,000DC-101 Decanter Centrifuge 1 40.00 m3/h SS316 293,000SL-102 Silo/Bin 1 75.00 m3 Concrete 75,000TFE-101 Thin Film Evaporator 1 20.00 m2 SS316 499,000V-108 Receiver Tank 1 3.80 m3 CS 20,000HX-107 Heat Exchanger 1 37.44 m2 CS 72,000V-107 Blending Tank 1 10.00 m3 SS316 265,000SL-103 Silo/Bin 1 2.00 m3 CS 75,000

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V-111 Receiver Tank 1

PM-110 Centrifugal Pump 1

HX-108 Heat Exchanger 1

PC-101 Pneumatic Conveyor 1

ant PUFA for those applications are �-linolenic acid (C18:3n–3,LA), �-linolenic acid (C18:3n–6, GLA); eicosapentaenoic acid

C20:5n–3; EPA), and docosahexaenoic acid (C22:6 n–3, DHA)hat are becoming high demanded in food and pharmaceuticalndustries. Heterotrophic fungi (Mortierella, Pythium, Mucor), andacteria (Shewanella, Moritella) produce EPA-rich oil in relatively

arge amounts from simple carbon. Marine organisms, such asinoflagellates and traustochytrid protists, such as Crypthecodiniumohnii and Traustochytrium accumulates more than 50% of theiripids as DHA using submerged cultivation in bioreactor being

mportant sources of industrial PUFA producers (Bellou et al., 2016;atledge, 2013).

SCO production of lipids is regarded as a partially growth-ssociated bioprocess. In general, it is explored in the literature as

0 m3 CS 20,0008 kW SS316 10,0000 m2 CS 32,0000 m CS 97,000

a two phases bioprocess: a growth phase followed by a lipid pro-duction (accumulation) phase (Ratledge and Cohen, 2008). Duringgrowth phase microorganism is cultivated in a well balance cul-ture media aiming to the cell proliferation at high specific growthrate to obtain a high-cell concentration. The accumulation phaseis achieved afterwards due to a limited nutrient (nitrogen, phos-phorus and/or sulfates) concentration and a non-limited carbonsource policy (Papanikolaou et al., 2009). The high carbon-nutrientratio triggers intracellular lipid accumulation from acetyl-CoA thatis directed to the de novo lipid biosynthesis (Ratledge, 2014). Liter-

ature also reported the limitation of iron and oxygen as affectingthe lipid accumulation (Granger et al., 1993; Hassan et al., 1993).The carbon source and investment costs in bioreactor are of greatimpact in SCO production cost. Therefore, high-productivity and

4 ops and Products 89 (2016) 478–485

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Table 3Direct Fixed Capital Cost of microbial lipid production (considering 316L stainlesssteel bioreactor).

Parameters Value (US$) % Investment Factors

Equipment Purchase Cost (EPC) 39,442,000 44.47% 1.00 EPCInstallation 12,205,000 13.76% 0.31 EPCProcess Piping 2,367,000 2.67% 0.06 EPCInstrumentation 3,944,000 4.45% 0.10 EPCInsulation 394,000 0.44% 0.01 EPCElectrical 789,000 0.89% 0.02 EPCBuildings 2,367,000 2.67% 0.06 EPCYard Improvement 394,000 0.44% 0.01 EPCAuxiliary Facilities 2,367,000 2.67% 0.06 EPCEngineering 6,427,000 7.25% 0.16 EPCConstruction 6,427,000 7.25% 0.16 EPCContractor’s Fee 3,856,000 4.35% 0.10 EPCContingency 7,712,000 8.70% 0.20 EPCDirect Fixed Capital Cost 88,691,000 100.00% 2.25 EPC

Table 4Operating cost of microbial lipid and deffated biomass production on industrialplant.

Cost Item Amount Unit $ %

Raw Materials 13,314,000 46.01Molasses 123,557 ton 12,355,694Amm. Sulfate 9,586,721 kg 958,672Water 330,174,407 kg 0Air 3,212,261,717 kg 0

Labor-Dependent 2,122,000 7.33Facility-Dependent 5,744,000 19.85Utilities 7,759,000 26.81

Std Power 194,548,440 kW-h 5,836,453

80 J.P.F. Vieira et al. / Industrial Cr

igh-lipid concentration and the use of low-cost carbon sources,uch as agricultural residues are pivotal to achieve lipid biopro-ess production at a competitive price (Ageitos et al., 2011). In thisontext several attempts has being carried out to develop a higherformance bioprocess using a diversity of carbon sources. Thewo-phases (growth phase followed by accumulation phase) fed-atch process carried out for several investigators in the past years

s an attractive procedure (Li et al., 2007; Zhu et al., 2008; Zhao et al.,012; Munch et al., 2015). It is flexible to control the growth and

ipid accumulation phases by modifying the feed flowrate of carbonource and limiting nutrient throughout the fermentation processnd is largely applied in bioprocessing to achieve high-cell densityigh-productivity production of biomolecules. In theory, duringhe late stage of the growth phase, nitrogen (and/or phospho-us) sources exhaustion (or limitation) impairs the cellular growth,hich leads to a reduction in the cell mass production rate and the

hanneling of the flux of carbon toward lipid biosynthesis. Recentlyell-succeeded attempts to produce lipids using solid-state fer-entation directly from cellulosic indicated that this alternative

oute might be an interesting option to decrease lipid productionost (Cheirsilp and Kitcha, 2015).

Sugarcane (Saccharum sp) molasses is a raw material rich inalts and carbohydrates (sucrose, glucose and fructose). Molassess a residue from sugar (sucrose) production from sugarcane and isroduced at a rate of about 40–60 kg of molasses/ton of processedugarcane. (Meade and Chen, 1977). We recently demonstrated theotential of lipid production from sugarcane molasses using theleaginous yeasts Rhodotorula glutinis, Rhodosporidium toruloides,hodotorula minuta and Lipomyces starkey (Vieira et al., 2014).aking into account the amount of sugarcane devoted to sugarroduction, that was about 280 mton harvested sugarcane in the015/2016 season (CONAB, 2015), molasses production in Brazilas estimated to be 13.000.000 t/year, with carbohydrate content

f about 85% m/m. This carbohydrate therefore is a potential carbonource to be used for SCO production. Moreover, the integration ofCO production in a sugarcane-mill would help to decrease lipidost because of ease of availability of raw material, utilities (steamnd electricity), effluent treatment and disposal and logistic distri-ution. This scheme was proposed in the past in order to produceacterial poly-hydroxyalkanote biodegradable polymer from sug-rcane carbohydrates (Rossell et al., 2001; Bueno Netto et al., 2000).

In this context, this study aimed to develop a process designo produce lipid using sugarcane molasses at high-cell-densitynd high-lipid-cell content with a selected oleaginous yeast. Therocess was integrated in silico on a typical sugarcane mill Brazil-

an industry, and accessed its economic viability based on kineticesults for different scenarios.

. Material and methods

.1. Microorganisms and culture media

R. glutinis CCT 2182 was obtained from the Andre Tozellooundation Culture Collection (Campinas, SP, Brazil) and it was pre-iously selected as detailed elsewhere (Vieira et al., 2014). The firsteed culture medium was YPD, composed of yeast extract (3.0 g l−1),eptone (5.0 g l−1), and glucose (10 g l−1). The second seed culturemployed a modified mineral medium containing (g l−1): glu-ose (20.0), (NH4)2SO4 (1.0), KH2PO4 (1.0), Na2HPO4·12 H2O (1.0),g2SO4·7H2O (2.0), NaCl (1.0), CaCl2 (0.02), FeCl3·6H2O (0.01) and

east extract (2.0). The cultures were incubated in a shaker (Excella24, New Brunswick Scientific, Edison, USA) for 24 h at 250 rpm,8 ◦C, and pH 5.8, and were then employed as inoculum in theioreactor cultures. The main substrate used in the bioreactor cul-ure media was Brazilian molasses (obtained from Usina da Pedra,

Steam 898,276 ton 89,828Cooling Water 26,521,961 ton 1,326,098Chilled Water 5,070,426 ton 507,043

Total 28,940,000 100.00

São Paulo, Brazil), composed of sucrose (356.87 ± 9.08 g l−1), glu-cose (152.35 ± 0.55 g l−1), fructose (121.30 ± 1.21 g l−1), and NH4

+

(0.5 ± 0.05 g l−1), with a C/N ratio of 40.

2.2. Fed strategy

In order to obtain high-cell-density with high lipid content,a carbon non-limited fed batch feeding strategy was used anddescribed as following.

A suspension of R. glutinis CCT 2182 previously prepared asin item 2.1 was inoculated at 10% of working volume in a benchbioreactor (Bioflo 115, New Brunswick Scientific, Edison, USA)with 2 l working volume. The bioreactor has an initial total reduc-ing sugar (TRS) concentration from molasses 18.0 g l−1, (NH4)2SO42.3 g l−1 supplemented with 0.80999 g KH2PO4; 0.73636 g Mg2SO4and 0.73636 g yeast extract per 10 g of total reducing sugar. Soonafter inoculation, an exponential fed of sugarcane molasses solution(SCMS) at 300 g TRS l−1 was set up in order to keep TRS concentra-tion in fermentation broth in a non-limited concentration (in therange of 10–20 g l−1), according to Eq. (1)

F = �xXoVo

(So − S) YX/Se�x�t (1)

where, F is the volumetric flow rate of SCMS varied with time �t,�x is the specific growth rate, Yx/s is the cell mass yield factor fromconsumed TRS, Vo is the initial volume of fermentation broth inbioreactor, Xo is the initial cell concentration and So is the TRS con-

centration in the feed stream. The adopted parameters values forEq. (1) were: So = 300 g l−1 �x = 0.23 h−1, Yx/s = 0.53 g biomass g−1

TRS; Vo = 1.5 l; and Xo = 1.00 g l−1.In this phase the pH 4.8 was controlled by a solution of

ammonium hydroxide (6.25% by volume) in order to provide a

J.P.F. Vieira et al. / Industrial Crops an

Table 5Economic evaluation considering 316L stainless steel bioreactor.

Parameters Value Unit

Annual operating time 8.300 hTotal Capital Investment 95,137,000 $

Direct Fixed Capital 88,691,000 $Working Capital 2,012,000 $Startup Cost 4,435,000 $

Capital Investment Charged to This Project 95,137,000 $Operating Cost 28,940,000 $/yrMain Revenue 22,572,000 $/yrOther Revenues 10,804,339 $/yrTotal Revenues 33,376,000 $/yrCost Basis Annual Rate 16,719.84 ton MP*/yrUnit Production Cost 1,730.86 $/ton MP*Unit Production Revenue 1,996.20 $/ton MP*Gross Margin 13.29 %Return On Investment 7.62 %Payback Time 13.13 years

*

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IRR (Before Taxes) 8.05 %NPV (at 7.0% Interest) 8,323,000 $

MP: Main product.

on-limited nitrogen regime. As soon as the yeast cell concentra-ion reached about 20.0 g l−1, the pH control solution was changedrom ammonium hydroxide solution to NaOH solution 3 M, in ordero impose a nitrogen-limited regime, triggering therefore lipidccumulation. Temperature, pH, aeration, and agitation rate wereutomatically controlled at 28 ◦C and 4.8. respectively. The aerationnd agitation rate were automatically controlled in the interval of.5–2 vvm (1.25–5.0 l of air min−1), and 250–1200 rpm to main-ain dissolved oxygen concentration in fermentation broth (D.O.)t 20% of air saturation. The experiments were carried in triplicatend results are presented as mean values and standard deviation.

.3. Chemical analyses

Sampling was withdraw periodically, centrifuged at 10.000g,0 ◦C, and chemical analyses performed in supernatant and yeastell sediment. The concentrations of glucose, fructose, and sucroseere measured in the supernatant using high performance liquid

hromatography (HPLC) Ultimate 3000, Dionex. The total concen-ration of reducing sugars (TRS) was calculated as the sum of theoncentrations of the individual sugars. Ammonia nitrogen con-entration was determined in the supernatant accordingly (Srienct al., 1984); the total lipid concentration was determined gravi-etrically and calorimetrically after lysis of the yeast (Fales, 1971;

ligh and Dyer, 1959). Disruption of the yeast cells and lipid recov-ry was performed as described (Vieira et al., 2014). Methyl esters ofatty acids were obtained by direct transesterification (Lewis et al.,000) and determined by gas chromatography. Detailing of appliedethodologies were described elsewhere (Vieira et al., 2014).

.4. Economic analysis hypothesis

Microbial lipid production and defatted biomass plant produc-ion was proposed at a nominal capacity of respectively 16,720 tonnd 21,600 ton per year, running 8300 h per year. The designedndustrial plant is supposed to sell the produced lipid and the defat-ed yeast as animal food supplement. The projected selling prices ofhese products were US$ 1,300.00/ton of lipid and US$ 500.00/tonf defatted yeast. Labor, molasses, and nitrogen were respectivelyS$ 5.29/h of labor, US$ 100.00/ton of molasses raw material and

S$ 0.100/ton of nitrogen source. Purchase prices of water, satu-

ated steam and electricity were set at US$ 0.050/ton water, US$.100/ton steam and US$ 0.030/kW h. The project design economicvaluation used the following hypothesis: (i) construction periodf 12 months; (ii) startup period of 6 months; (iii) project lifetime

d Products 89 (2016) 478–485 481

of 35 years; (iv) NPV interest rate of 7%; 10% and 12%; (v) 100% ofoutlay of investment in the 1st year; (vi) depreciation straight lineover 30 year; (vii) salvage value of direct fixed capital of 5%; (viii)maintenance of 1.7% of direct fixed capital per year. (ix) insuranceof 0.5 % of direct fixed capital per year; (x) taxes of 1.0 % of directfixed capital per year; (xi) factory expense of 1.0 % of direct fixedcapital per year; (xii) working capital of 30.0 % of raw material;utilities and labor; (xiii) startup cost of 5.0% of direct fixed capi-tal; (xiv) selling taxes of 0% of gross profit. SuperPro Designer v8(Intelligen, Inc., USA) was used for estimation of fixed capital andoperating costs. Required power for agitation in fermenter Pf (kW)was estimated by Eq. (1) (Koutinas et al., 2014).

2.5. Lipid stoichiometry equation

The adopted stoichiometry of lipid and yeast production fromglucose (Eqs. (2) and (3)) was described elsewhere (Koutinas et al.,2014):

C6H12O6 + 4.32O2 + 0.54C5.35H9.85O2.45N1.5

→ 1.12C4H6.5O1.9N0.7 + 4.41CO2 + 5.02H2O (2)

C6H12O6 + 1.26O2 → 0.06C57H104O6 + 2.61CO2 + 2.92H2O (3)

Eq. (4) shows a modification of Eq. (2) for yeast biomass pro-duction from the ammonium ion present in the ammonium sulfate(NH3

+) and glucose (stoichiometric balance).

C6H12O6 + 2.87O2 + 0.26(NH4)2SO4

→ 0.76C4H6.5O1.9N0.7 + 2.97CO2 + 4.60H2O + 0.26SO4−2 (4)

The theoretical mass conversion yield of glucose and nitrogenfrom ammonium sulfate to yeast biomass are equal to 0.40, and9.67, respectively. The theoretical mass conversion yield of glucoseand nitrogen from yeast extract to yeast biomass and lipids areequal to 0.589, 9.35 and 0.295, respectively. According to the resultsobtained in this study, the conversions of glucose to yeast biomassand glucose to nitrogen are equal to 0.42, and 8.71, respectively.The conversion values were intermediate between cell growth withammonium sulphate and cell growth with yeast extract due to sup-plementation of the medium with a mix of these nitrogen sources.However, the stoichiometric equation by Koutinas et al. (2014)did not represent the lipid fraction of the yeast biomass, whichis approximately a lipid mass fraction of 7% (m/m). Then one canrearrange these equations for the cell mass production accordingto Eqs. (5) and (6), using respectively ammonium sulfate and yeastextract.

C6H12O6 + 2.87O2 + 0.26(NH4)2SO4 → 0.76C3.57H5.72O1.86N0.7

+2.97CO2 + 4.60H2O + 0.01C57H104O6 + 0.26SO4−2 (5)

C6H12O6 + 4.23O2 + 0.54C5.35H9.85O2.45N1.5

→ 1.12C3.71H1.70O1.86N0.72 + 4.41CO2 + 5.02H2O

+0.01C57H104O6 (6)

2.6. Power consumption

The required power for agitation in fermenter Pf (kW) was esti-mated by Eq. (7) (Koutinas et al., 2014). Power consumption of other

482 J.P.F. Vieira et al. / Industrial Crops and Products 89 (2016) 478–485

F the no

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dia

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ig. 1. Time evolution for Rhodotorula glutinis CCT 2182 bioreactor cultivation in

oncentration; (�) TRS concentration; ( ) nitrogen concentration ( ) total yeand standard deviation of triplicate experiments.

evices was estimated with the aid of Superpro® Designer accord-ng to the flow rate, density of the fluid estimated pressure pipelinend pressure on the equipment (see Table 5 ).

f = 1�˛

2.8�Vf (7)

here:�� = electrical motor efficiency (0.9)�= Fraction workingolume/size of fermenterVf = Fermenters working volume (m3)

. Results and discussion

.1. Bioreactor strategy for lipid production

The present study proposed a SCO production approach thatmulated bacterial polyhydroxyalkanoates cytoplasm accumula-ion applied in past studies (Diniz et al., 2004; Pradella et al., 2010).he protocol comprises the fed-batch high-cell-density growthhase followed by a fed-batch nitrogen-limited lipid accumula-ion phase (Fig. 1a). R. glutinis grown attained total yeast biomassoncentration (X), maximum specific growth rate (�x) lipid pro-uctivity (Pr) and lipid yield from TRS respectively of 62.25 g l−1,.235 h−1; 0.425 g l−1 h−1 and 0.171 g of lipid g−1 of TRS. Theell mass kinetic displayed an exponential profile until approx-mately 24 h, indicating a balanced growth for proposed culture

edium and conditions. Due to nitrogen limitation after 24 h,efatted biomass concentration (Xr) remained constant at approx-

mately 30 g l−1. Lipid accumulated in yeast cell climbed to 54%f dry cell weight at the end of experiment. A comparison ofhe kinetic data obtained in this study with the results obtainedy other researchers that used R. glutinis (or its anamorph formhodosporodium toruloides) aiming the production of lipids from

diversity of carbohydrates was carried out (Table 1). Reportedipid production from glucose attained higher values of cell con-entration and lipid productivity than obtained results and similarmount of cell-lipid content (Table 1). Although we tried mainte-

ance of D.O. above 20% of air saturation during the experiments,his was not achieved between 40 h to 60 h fermentation time. Inhis period, the D.O. went down to less than 2% while aeration andgitation rate climb at its maximum levels, respectively, 5.0 l ofir min−1 and 1200 rpm (data not shown). We want to emphasize

n-carbon limited fed strategy and supplemented culture media: ( ) total lipids

ass concentration; ( ) defatted biomass concentration. Points are mean values

that in the proposed protocol the broth dissolved oxygen concen-tration was automatically controlled by variation of aeration andagitation rate. Cultivation in previous studies (Liu et al., 2015; Zhaoet al., 2012) used oxygen-air enrichment in order to supply the highoxygen demand due to high-cell density cultivation. These strate-gies however would pose an undesirable cost increment in the SCOproduction and generally are only justified for high-value products.

Lignocellulosic hydrolysate composed a mixture of glucose andxylose, submitted to detoxification and supplemented with saltsolution was also recently explored as alternative carbon sourcein order to decrease the impact of raw material cost. However, verylow Pr was obtained (about 0.2 g l−1 h) (Liu et al., 2015; Anschauet al., 2014) (Table 1). This would probably hamper its commercialviability due to the high investment costs in bioreactor. The resultspresented by this study led to lower results in terms of cell densityand cell-lipid content but have the merit of the usage of sugarcanemolasses, a carbon low-cost agricultural residue commercialized atabout US$ 100/ton in this country.

3.2. Proposed process flow diagram

Exploitation of microorganisms as a source of bioenergy pro-duction could meet technical and economic conditions as biofuelresource. Some advantages has being raised in literature as costcompetitiveness with petroleum fuels, air quality improvement(e.g. CO2 sequestration), and minimal water usage requirements(Tabatabaei et al., 2011). Some studies of new technologies inte-grated in sugar mills has been done recently (Dias et al., 2011;Furlan et al., 2013). Residual sugars such as molasses obtained fromthe sucrose commercial sugar refining is a potential carbon sourcematerial candidate as it is composed mainly by sucrose, glucose andfructose, salts and vitamins (Hugot, 1969; Meade and Chen, 1977).Its usage was recently proposed as raw material for lipid productionutilizing ligneous yeasts (Vieira et al., 2014).

A Process Flow Diagram (PFD) for microbial lipid production

from molasses was therefore proposed based on the developed pro-tocols and includes the main streams and equipment involved inthis process (Fig. 2). The stream components and its specificationare shown (Annex A, Table A3). In short, culture medium composedof sugarcane molasses supplemented (MO001/002 + NS001/002)

J.P.F. Vieira et al. / Industrial Crops and Products 89 (2016) 478–485 483

obial

aaC1fd

Fig. 2. Process flow diagram (PFD) for micr

re pumping to heat exchangers (HX-101 and HX-102) to attaindequate disinfection level and received in tanks V-103 and V-106.

ulture media ME-101 from V-103 is supplied to bioreactors R-01 and R-102 to produce yeast inoculum. Culture media ME-102rom V-106 is supplied to R-103 bioreactors for lipid process pro-uction. Yeast oleaginous suspension (YO-001) from bioreactor is

lipid production from sugarcane molasses.

centrifuged in CF-101 and the produced yeast cream is storage inreceiver tank V-109. Yeast concentration and drying is carried out

in flash drum (V-110) followed by spray drying process (SDR-101).The dried yeast Y-003 is suspended in hexane HX-001 in the blend-ing tank V-107 and disrupted in the high-pressure cell homogenizerHG-101. Defatted biomass (RYD-001) is separated from hexane oil

4 ops and Products 89 (2016) 478–485

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Table 6Economic evaluation considering epoxy lined carbon steel bioreactor.

Parameters Value Unit

Annual operating time 8.300 hTotal Capital Investment 47,022,000 $

Direct Fixed Capital 42,867,000 $Working Capital 2,012,000 $Startup Cost 2,143,000 $

Capital Investment Charged to This Project 47,022,000 $Operating Cost 25,972,000 $/yr

Raw Materials 13,314,000 $Labor-Dependent 2,122,000 $Facility-Dependent 2,776,000 $Utilities 7,759,000 $

Main Revenue 22,572,000 $/yrOther Revenues 10,804,339 $/yrTotal Revenues 33,376,000 $/yrCost Basis Annual Rate 16,719.84 ton MP/yrUnit Production Cost 1,553.36 $/ton MPUnit Production Revenue 1,996.20 $/ton MPGross Margin 22.18 %Return On Investment 18.63 %Payback Time 5.37 yearsIRR (Before Taxes) 24.61 %

84 J.P.F. Vieira et al. / Industrial Cr

olution (OS-001) in decanter DC-101, cleaned from solvent (nothown) to render 2.60 ton h−1 of product. Microbial lipids (SCO-04) are recovered from the hexane oil solution (OS-001) in thehin film evaporator TFE-101 to render of 2.01 ton h−1 of product.

.3. Economic analysis

The PFD summarized in Fig. 2 possess a main inlet streamf molasses that is bioprocess carbon source acquired at US$00.00/ton and two main outlets microbial lipid and defattediomass with a selling price fixed at respectively US$ 1,300.00/tonf lipid and US$ 500.00/ton of defatted yeast. The microbial lipidalue was set at US$ 1330/ton as an intermediary value of theegetable oil practiced in US during 2003–2013 periods (Koutinast al., 2014). The defatted biomass value was set at US$ 500/ton tomulate soybean meal quotation in the same period.

The list of equipment specifications and unitary cost of acquisi-ion of PFD (Fig. 2) is presented in Table 2.

Bioreactor train is by far the most expensive item of investmentost accounting followed by separation equipment being thereforeipid productivity and lipid concentration two of the most impor-ant key process variable to be optimized.

Based on mass balance and the cost of acquisition of equipmentor the process described in Fig. 2 the Direct Fixed Capital Cost isresented in Table 3. It is worth to mention that in the presentlternative it was choose to use full-sterilized stainless steel 316Lioreactors conventionally used the bio industry, and as expectedhe DFCC climb to a value very high of US$ 88.691.000.

Operating cost for the proposed flow diagram was therefore cal-ulated taking into account the annual production of microbial oilnd defatted biomass. Table 4 shows the operating costs of micro-ial lipid production of the proposed industrial plant.

Based on Tables 3 and 4 a profitability economic analysis of theroposed industrial plant for the 316L stainless steel bioreactorption was carried out and summarized in Table 5.

The analysis indicated that the lipid plus defatted yeast produc-ion based in the present data was a marginally economic feasibleperation. The IRR attained 7.62% per year and the NPV value (at 7%RR per year) was about US$ 8,323,000 (before taxes) for molassesrice at US$ 100.00/ton and the selling prices of microbial lipid andefatted biomass respectively at US$ 1,300.00/ton of lipid and US$00.00/ton of defatted yeast.

Ratledge and Cohen (2008) reported an expected price of micro-ial oil to be US$ 3000/ton in order to make the process economicttractive, not considering the cost of the carbon source. Therefore,he selling price would expect to be higher than this value if it wouldonsider the carbon source cost. This is in agreement with Koutinast al. (2014) that reported a minimum lipid oil price of about US$000/ton at a carbon source cost of US$ 100/ton, (the same valuesed in the present study). Even so, this value would not competeith the vegetable oil price that varied between US$ 1000 to US$

000/ton during 2003–2013 periods (Koutinas et al., 2014).In present analysis the Stainless steel 316L stirred tank bioreac-

or (R-102 and R-103, Table 2) is by far the most expensive item ofapital investment in the production plant representing about 90%f equipment purchase cost hence impacting total investment costTable 3) and making a very tight economic operation. Therefore,t is of the most importance to have alternative bioreactor equip-

ent in order to alleviate the investment cost for SCO production.poxy-lined carbon steel bioreactor is a commonly used equipment

n ethanol production in Brazil (JP Vieira, personal communication).stimative of capital, operating costs, as well the main parametersf the process economic evaluation for this alternative configura-ion was therefore carried out (Table 6). The economic analysesere made on the same proposed PFD (Fig. 2).

NPV (at 7.0% Interest) 67,797,000 $

*MP: Main product.

The use of this equipment would bring the investment cost fromabout US$ 88,691,000 (Table 5) down to US$ 42,867,000 (Table 6).

Consequently, in this alternative the internal rate of return (IRR)would climb up to 24.61% leading the NPV (at 7% interest rate) ofUS$ 67,797,000/year (before taxes). Therefore, the proposed mod-ification brings the process potentially attractive.

The increase of lipid productivity (mass of lipid/volume of biore-actor × hour) and the use of low-cost carbon source are obviouschoices to decrease SCO cost production. Coproducts such as defat-ted biomass selling at US$ 500/ton considered in the present studywould help to benefit industrial plant cash flow. Another alternativepreclude in literature is the use of microbes (naturally or molecularbiologically constructed) able to store high-values fatty acids, suchas the �-3 fatty acids, to help increase industrial plant cash flow(Ageitos et al., 2011).

However, we think that the use of low-cost bioreactor as pro-posed in the present study appears to be a simple alternative withinteresting potential deserving attention to be explored in thefuture.

4. Conclusions

The growth and lipid accumulation phases of R. glutinis on sug-arcane molasses as raw material. was studied in bioreactor andbest attained typically total biomass concentration 62.25 g l−1, lipidproductivity 0.42 g l−1 h−1, sugar conversion yield to lipids 0.21 gof lipid g−1of total reducing sugar. An industrial bioprocess con-ceptual design including mass balance, list of equipment, cost andspecifications was established. Economic performed analysis indi-cated the use of low cost bioreactor built in carbon steel coatedwith epoxy resin to reduce CAPEX for a substantial increase in NPVto make project be economically attractive.

Acknowledgements

The present project was partially financed by FAPESP. The

authors want to thanks to the Brazilian Bioethanol Science andTechnology Laboratory (CTBE) of National Center of Energy andMaterial (CNPEM) for the use of its installation and the technicalassistance.

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growth and lipid accumulation of oleaginous yeast Rhodosporidium toruloidesand preparation of biodiesel by enzymatic transesterification of the lipid.

J.P.F. Vieira et al. / Industrial Cr

ppendix A. Supplementary data

Supplementary data associated with this article can be found, inhe online version, at http://dx.doi.org/10.1016/j.indcrop.2016.05.46.

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