Double-effect distillation and thermal integration applied to theethanol production process
Reynaldo Palacios-Bereche a, *, Adriano V. Ensinas a, b, Marcelo Modesto a,Silvia A. Nebra a, c
a Centre of Engineering, Modelling and Social Sciences (CECS/UFABC), Federal University of ABC, Av. dos Estados, 5001, CEP 09210-580, Santo Andr�e, SP,Brazilb �Ecole Polytechnique F�ed�erale de Lausanne- STI-IGM-IPESE, Station 9, 1015 Lausanne, Switzerlandc Interdisciplinary Centre of Energy Planning (NIPE/UNICAMP), University of Campinas, Rua Cora Coralina, 330, CEP 13083-896, Campinas, SP, Brazil
a r t i c l e i n f o
Article history:Received 2 July 2014Received in revised form2 December 2014Accepted 20 January 2015Available online 11 February 2015
A double-effect distillation system allows a significant reduction in energy consumption, since thecondensers and reboilers of different columns can be integrated thermally. To achieve this goal, somecolumns operate under a vacuum, while others operate close to atmospheric pressure. These pressurelevels bring about different temperature levels, allowing energy recovery. Thus, the aim of this study is toassess the incorporation of double-effect distillation in ethanol production, and its impact on energyconsumption and electricity surplus production in the cogeneration system. Moreover, because double-effect distillation and thermal integration involve an increase in equipment costs, an economic assess-ment was done. Several cases were evaluated and a thermal integration technique was applied, in orderto integrate the overall process. The thermal integration study showed that it is possible to integrate thejuice concentration step (multiple effect evaporation system) in the overall process without additionalthermal consumption, through the selection of a suitable set of pressures in the evaporation system. Theresults showed a reduction in steam consumption of between 17% and 54%, in comparison with the BaseCase. Regarding the electricity surplus, this increased by up to 22% when extractionecondensing steamturbines were adopted.
The multiple-effect operation of distillation and rectificationcolumns allows a significant reduction in energy consumption,since the condensers and reboilers of different columns can beintegrated thermally [1].
Dias et al. [2] studied the effect of double-effect distillationprocess, using the Hysys® simulator. The configuration studied bythese authors is similar to the conventional distillation used inethanol production plants from sugarcane, but with the distillationcolumns operating under a vacuum (20e25 kPa), while the recti-fying columns operate close to atmospheric pressure (101 kPa atthe top). According to these authors, the temperature difference
between the distillation column reboiler and the condensers of therectifying and dehydration columns (these authors assumed adehydration extractive column with monoethylene glycol; MEG),allows the integration of these devices, reducing the steam con-sumption in the distillation process significantly.
Junqueira et al. [1] accomplished a simulation study using theAspen Plus® simulator, to evaluate alternative configurations toconventional fermentation and distillation in the ethanol produc-tion process. These alternatives included vacuum extractivefermentation coupled with double-effect distillation and triple-effect distillation. Triple-effect distillation is another configurationof thermal integration, where distillation columns operate under avacuum (19e25 kPa), and liquid phlegm, produced in the top col-umn of a distillation set, is divided into two streams: the first is fedinto a rectification column that operates close to atmosphericpressure (70e80 kPa); the other is fed into a rectification columnoperating under a relatively high pressure (240e250 kPa). Ac-cording to Junqueira et al. [1], the vacuum extractive fermentation
C Monetary cost of a stream (US$/h)c Monetary cost per unit of exergy (US$/kJ)ccw Cooling water cost (US$/m3)CC Composite curvesE Equipment cost (US$)ex Specific exergy (kJ/kg)GCC Grand composite curveHTST High temperature short timei Annual interest ratej Equipment useful lifeLHV Lowing heating value_m Mass flow (kg/h)MEG Monoethylene glycol
MEE Multiple-effect evaporatorP Annual cost (US$/year)t Factory operation hours per yearT* Shifted temperaturev Specific volume (m3/kg)_Z Investment cost (US$/h)
Subscriptss Steambag Bagassew Waterb Boilercw Cooling water
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coupled with triple-effect distillation presented lower energyconsumption (reduction in steam consumption by 36% in com-parison to the conventional process).
Bessa et al. [3] studied distillation columns integrated thermallyfor ethanol production, considering a large amount of minorcompounds. The double-effect distillation system studied by theseauthors was composed of two sets of columns. The first set oper-ated at a pressure greater than atmospheric pressure (1.52 bar) andthe second one, under a vacuum (0.219 bar).
Siemens Ltd and Dedini S/A developed a system called a Split-Feed (patent PI0600553 BR-5), based on the multiple-effect distil-lation column principle. In this system, pressurized columns areresponsible for generating sufficient heat to operate the columnsunder a vacuum [4]. The authors recommend that the heatexchanger reboiler-condenser should be of a falling film type, sincethis type of heat exchanger has a high heat transfer coefficient andcan operate with small temperature differences.
Thus, based on previous research, the aim of this study is toassess the incorporation of double-effect distillation in ethanolproduction, and its impact on the energy consumption and elec-tricity surplus production in the cogeneration system. Moreover,since double-effect distillation and thermal integration cause aninherent increase in equipment costs an economic assessment isdone. Simulations in Aspen Plus® software were performed, in or-der to evaluate the mass and energy balances. Several cases wereevaluated and a thermal integration technique applied, in order tointegrate the overall process.
Simulation of the process was carried out using Aspen Plus®
software, according to [5] and [6]. The model selected for propertycalculations in the simulator depended on the operation type. Inthe cogeneration system, the Redlich-Kwong-Soave equation withthe Boston Mathias function was adopted for combustion gases,while for water, the Steam Tables method of the Aspen Plus® wasused. For the sucrose-water solution (sugarcane juice), the UNI-QUAC model was selected, with the binary parameters of Starzacand Malthouthi [7]. For the enthalpy calculation of juice, a sub-routine was written in Fortran, according to equations from theliterature [8]. Finally, for ethanol mixtures in the distillation anddehydration steps, the UNIQUACmodel was selected. Regarding thecomponents adopted in simulation, some constituents of sugarcaneare not found in the Aspen Plus® database, thus they were created,and their properties inserted into the software, according to data
from the literature [9]. Cellulose, hemicellulose and lignin wereselected to represent the fibre in sugarcane. Reducing sugars weresimulated as glucose, minerals as K2O, and other non-saccharidesas aconitic acid and potassium chloride. The soil in sugarcane wassimulated as SiO2 [5,10]. Table 1 shows the components adopted inthis study.
2.1. Modelling and simulation of the ethanol production process
Fig. 1 presents a simplified block diagram of the ethanol pro-duction process adopted in this study.
To begin the modelling it was necessary to define the compo-sition of sugarcane that arrives at the factory. Table 2 shows thesugarcane composition adopted in this study.
Fig. 2 shows a detailed flow sheet of the conventional produc-tion process. It was divided into the following subsystems:
2.1.1. Sugarcane cleaning, preparation and juice extractionThe cleaning system is used to remove field soil, rocks and
rubbish that came with the sugarcane. After cleaning, sugarcane isprepared using rotary knives and shredders, in order to reduce thesize of the pieces, as well as to rupture the sugar-bearing cells in thecane to facilitate the extraction of the sugar process [11]. The juiceextraction separates the sucrose-containing juice from the bagasse.Bagasse is used as fuel in a cogeneration system, while the juice is
Fig. 1. Simplified block flow of the ethanol production process from sugarcane.
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sent to the treatment process. In this simulation, cane dry cleaningand a mill tandem extraction system were assumed.
2.1.2. Juice treatmentThe juice from extraction system contains a significant quantity
of small pieces of bagasse, suspended matter and non-sugar im-purities. Thus, in juice treatment several operations are adopted inorder to remove these impurities. In this simulation, the followingoperations were adopted: screening, heating, liming, flashing anddecantation. The clarified juice obtained in a decanter was sent tothe concentration step, while the precipitated was sent to the mudfiltration. The filtrate was recycled to the heating operation whilethe solid residue (filter cake) was sent to the field [12].
2.1.3. Juice concentrationSince the must for ethanol production is prepared from sugar-
cane juice, treated juice should be concentrated to reach anappropriate sugar concentration for the fermentation process(approximately 17%). The concentration of treated juice takes placein a multiple-effect evaporation system (5 effects) [6,12]. Exhauststeam from the cogeneration systemwas used as an energy sourcefor the first effect. The multiple-effect evaporator works withdecreasing pressure, owing to a vacuum imposed on the last effect,producing the necessary temperature difference between theheating vapour and the juice along the consecutive effects. Vapourbleeds can be used to cover the heat requirements of other parts ofthe process [12]. In the conventional process, this study consideredthat vapour bleedings of the 1st effect are used for heating raw juicein the treatment step. According to Fig. 2, part of the juice goes
Table 2Sugarcane composition specified in simulation [6].
through the evaporation system, while the remainder by-passesconcentration and goes directly to must preparation.
2.1.4. Sterilization and coolingAfter the must preparation, there is the sterilization and cooling.
Must sterilization was carried out by a HTST-type treatment (HighTemperature Short Time), with heating to 130 �C, followed by fastcooling down to the fermentation temperature of 32 �C.
2.1.5. FermentationThe sugars of the must are converted in ethanol through a
biological process leads by yeast, Carbon dioxide is emitted in anexothermal reaction. In this study, fermentation was based on theMelle-Boinot process (cell-recycle batch fermentation). Afterfermentation the wine, containing approximately 6% of ethanol(mass basis), was taken to the distillation system to remove water.
2.1.6. DistillationThe wine is heated to a suitable temperature before entering the
first distillation column. Hydrated ethanol (93.7% wt. of ethanol) isobtained as a result of stripping and rectification stages. A largeamount of vinasse (bottom product of distillation column withapproximately 0.02% wt. of ethanol) is generated, which is handledas effluent. A detailed description of the distillation step is pre-sented in Section 2.2.
2.1.7. DehydrationIn order to remove the remaining water and produce anhydrous
ethanol, the dehydration stage is required. In this simulation, aprocess of extractive distillation with monoethylene glycol wassimulated. Two columns were adopted: the extractive column,EXTRAC; and the recovery column, RECUPERA (Fig. 2).
The main operational parameters of the modelled plant were:mill capacity, 2,000,000 t cane/year; crushing rate, 500 t cane/hour;season operations hours, 4000 h/year; and bagasse production,277 kg/t cane. Table 3 shows the equipment specifications for eachoperation of the conventional process.
Fig. 2. Flow sheet of conventional ethanol production process.
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2.2. Conventional distillation process
Fig. 3 presents the flow sheet of the conventional distillation inAspen Plus®. This configuration is based on the configurationcommonly employed in the industry, and it is composed of twosections: stripping of wine (comprised of three columns located ontop of each other: A, A1 and D); and stripping of phlegm (comprisedof two columns B and B1) [6,10]. Sections A1 and D are used mainlyfor reducing the contamination of ethanol with volatile com-pounds. Sections A and B1 are stripping sections, where ethanol isstripped from the liquid phase, obtaining the bottom productsvinasse and phlegmasse, respectively. These are two liquid streamsalmost free of ethanol [3]. Rectification takes place mainly in sec-tion B, concentrating ethanol in the vapour phase [3]. In this pro-cess, the wine is pre-heated in the condenser E and in the vinasseheat exchanger K (Fig. 3). After that, it is fed to the distillation set, inwhich vinasse (residue), phlegms (ethanol-rich streams, containingaround 40e50 wt% ethanol), second-grade ethanol (a mixture withhigh ethanol and volatiles content) and gases (mainly CO2) areseparated. Liquid phlegm is the bottom product of column D, andvapour phlegm is produced near the top of column A. Phlegms arefed to the rectification columns, producing hydrated ethanol in thetop and phlegmasse in the bottom.
2.3. Modelling and simulation of double-effect distillation system
The double-effect process is similar to the conventionalconfiguration (Fig. 4), but the distillation columns operate under a
vacuum, while rectification columns operate under atmosphericpressure. Thus, different temperature levels are observed betweenthe column A reboiler and the column BeB1 condenser, allowingthe thermal integration of these devices, and consequently,reducing energy consumption at the distillation stage. This inte-gration is done in the heat exchanger E�2 (Fig. 4).
Since there is a difference of pressure between distillation col-umns (D, A, A1) and the rectification column (BeB1), the liquidphlegm must be pumped to column BeB1, while the vapourphlegmmust be compressed in the unit COMP-S. In order to reducethe power consumption, a two-stage compression system withintercooling was assumed in the compression station COMP-S(Fig. 5).
Owing to the low pressure of column D, large amounts ofethanol are lost on the GS-1 stream, thus, an absorption column(washing column, COL-2) that uses a fraction of phlegmasse towash the gases, was adopted to recover some of the ethanolvaporized on top of column D, according to [1]. Another alternativeto reduce the ethanol lost in GS-1 stream could be to reduce thetemperature operation of condenser R. However, this option wouldrequire a refrigeration system, with its respective power con-sumption, to achieve a low enough refrigerant temperature.
In the proposal presented in Fig. 4, the flash tank FL-1 operatesat 25 kPa and the vapour phase (GS-1) is compressed (compressorCOMP D) and sent to the absorption column. The amount ofphlegmasse used in washing was adjusted, with the aim ofreducing the ethanol losses in the GASES stream, into the samerange as that in the conventional distillation. According to [13], in
Table 3Equipment specifications for each operation process [5,6].
Parameter Value
Sugarcane cleaning, preparation and juice extractionEfficiency of soil removal in cleaning operation, % 70Efficiency of sugar extraction in extraction system, % 96.2Imbibition water, kg/t cane 300Moisture content in bagasse, % 50Mineral content in raw juice, % 8.4Juice treatmentHeating temperature of juice treatment, �C 105Sucrose content in filter cake, % 2Moisture content in filter cake, % 70CaO consumption, kg/t cane 0.5Juice concentrationBrix content in final must, % 19Pressure 1st effect e evaporation system,a bar 1.69Pressure 2nd effect e evaporation system,a bar 1.31Pressure 3rd effect e evaporation system,a bar 0.93Pressure 4th effect e evaporation system,a bar 0.54Pressure 5th effect e evaporation system,a bar 0.16FermentationConversion yield from sugars to ethanol, % 89Fermentation temperature, �C 34Yeast concentration in fermentation reactor, v/v% 25Sulphuric acid for yeast treatment, kg/m3 of ethanol 5Distillation and rectificationNumber of stages in stripping section (column A) 18Number of stages in rectification section (column A1) 8Number of stages in top concentrator (column D) 6Number of stages in phlegm rectification column (column BeB1) 45Ethanol content in vinasse and phlegmasse, % 0.02DehydrationPressure in extractive column, bar 1.01Pressure in recovery column, bar 0.20Ethanol content in anhydrous ethanol, wt % 99.4
a Conventional production process (Base Case).
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conventional distillation, the ethanol losses in gas streams are nothigher than 0.25% of the ethanol that is fed to distillation columns.
Table 4 presents the main parameters for the simulation ofconventional and double-effect distillation.
2.4. Thermal integration
Although the double-effect distillation process represents athermal integration, this part of the study deals with the thermalintegration of the overall process. The Pinch Method was used toperform the thermal integration study [14e17]. The minimumapproach temperature difference (DTmin) adopted in this study forthe process streams was 10 �C, except for streams of the evapora-tion system, where 4 �C was selected. The existence of a multiple-effect evaporator (MEE) in the system represents a conceptualproblem for the construction of both Composite Curves (CCs) and aGrand Composite Curve (GCC), because the minimum target utilityis greatly affected by the arrangement and operation conditions.Moreover, vapour bleedings from MEE are used to fulfil heatingrequirements of different processes, and their flow rates are char-acteristic of these processes. To solve this issue, the thermal systemwas broken down into two subsystems: the MEE; and theremainder of the process. Therefore, the thermal integration pro-cedure was carried out applying the following iterative steps:
Step 1. Thermal integration of the available process streams, exceptin theevaporation system.Constructionof the initialGrandCompositeCurve (GCC). Table 5 shows the data of hot and cold streams availablefor thermal integration, without considering the MEE streams.
Step 2. Evaluate the possibility of appropriate placement for theMEE [18e20]. Since an evaporator is a thermal separator, each ef-fect of MEE can be represented as a box in a T*eH diagram, wherethe top side denotes the cold stream (juice) being heated, and the
bottom side denotes the hot stream (vapour) being cooled [20].Fig. 3a shows the GCC built in Step 1, and the evaporation effects asboxes on the left side of the graph. To achieve an appropriateplacement of MEE, the evaporator profile should not cross the GCC.However, the thermodynamic profile of a MEE can be manipulated,adopting appropriate operation pressures, increasing the numberof effects, or increasing the temperature difference across each box[14]. Moreover, the calculation of the appropriate vapour bleedingdemand in each effect is also possible using this method [18e20].
Step 3. Integration of the evaporation system and calculation ofthe energy targets (minimum requirements of hot and cold utility).
Step 4. Since vapour condensates are adopted as hot streams, it isnecessary to update of the mass rates of the evaporation systemstreams. Return to Step 2 until convergence.
2.5. Evaluated cases
In this study, three cases were evaluated and compared:Base: Ethanol production process with conventional distillation.DED: Ethanol production process with double-effect distillation
process.DEDTI: Ethanol production process with double-effect distilla-
tion and thermal integration of overall process.
2.6. Cogeneration system
Two cogeneration systems configurations were regarded in thisstudy:
Configuration I: The first configuration analysed is a steam cyclewith back-pressure steam turbines. In this case, the productionprocess determines the quantity of steam that can be produced bythe boiler, once there is no condensation system. On the other handthere is a bagasse surplus.
Configuration II: The second configuration is a steam cycle withextractionecondensing steam turbines. In this case, the condenseroffers more operation options and higher flexibility, making itpossible to operate all around the year [16]. The condensationpressure adopted was 0.1 bar.
Both configurations present a boiler with steam parameters of530 �C and 100 bar, and a juice extraction systemwith mills drivenby electric engines. Steam turbines have an extraction at 6 bar formust sterilization and ethanol dehydration, and at 2.5 bar for theother heating requirements of the process. The boiler and thecogeneration system were modelled according to previous studies[5] and [6], following what are regarded as the best practices in theindustrial sector. Fig. 6 shows the scheme of the cogenerationsystem for Configurations I and II. Table 6 presents the mainspecifications for the cogeneration system.
2.7. Economic assessment
While double-effect distillation and thermal integration pro-mote a reduction in utility consumption (steam and cooling water),their implementation involves additional investment. Thus, aneconomic assessment was accomplished in order to estimate theimpact of this additional investment on the process cost. Thisassessment includes the estimation of the equipment cost and theutility cost (heating and cooling) for the evaluated cases.
2.7.1. Equipment costFor the Base Case, the equipment cost was updated from the
literature data, based on Brazilian manufacturers [20,21].Since, in the double-effect distillation cases, columns A, A1 and D
operate under a vacuum, there is an inherent increase in equipmentsize and cost. Moreover, double-effect distillation requires additional
Fig. 3. Flow sheet of conventional distillation process.
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devices such as compressors and the absorption column (COL-2 inFig. 4).
The Aspen Process Economic Analyzer® is used to estimate thecost of devices required in the double-effect distillation system, andcorrelate them with the equipment costs in conventional distilla-tion (based on Aspen Plus database costs). From this relationshipand the data in Table 7, the equipment cost of a double-effectdistillation system is finally estimated.
With regard to the thermal integrated case DEDTI, the cost of theheat exchanger network was estimated, based on the heat transferarea and the reference costs of the Aspen Process EconomicAnalyzer®. Thus it was necessary to design the heat exchangernetwork that satisfies the targets of minimum energy consumption,to determine the new hot and cold utility consumption in each unit,as well as the mean temperature difference involved and the heattransfer coefficients. Moreover, since the operation conditionscould change in a multiple-effect evaporation system, the heattransfer areas, as well as their costs, were also recalculated.
2.7.2. Cost of utilitiesThe procedure proposed by Ensinas et al. [22] was applied, in
order to determine the steam cost. This procedure is based on theTheory of Exergetic Cost [23,24]. Thus the monetary cost Cs (US$/h)of each steam stream can be calculated using the following equation:
Cs ¼ cs _msexs (1)
where cs is the monetary cost per unit of exergy of steam (US$/kJ),_ms is the mass flow of steam (kg/h) and exs is the specific exergy ofsteam (kJ/kg).
The monetary cost per unit of exergy of steam generated inboiler is calculated by Eq. (2)
cs ¼�cbag _mbagexbag
�þ _Zb
_msðexs � exwÞ (2)
The monetary cost per unit of exergy of the bagasse (cbag), usedas fuel in the cogeneration system, was assumed to be the same asthe sugarcane that enters the factory (ccane), which can be calcu-lated from the production cost of sugarcane (Ccane), as follows:
ccane ¼ Ccanemcaneexcane
(3)
The investment cost of boiler _Zb (US$/h) can be calculated usingthe following equations:
_Zb ¼ Ebx (4)
x ¼ið1þiÞj
ð1þiÞj�1
t(5)
where Eb is the equipment cost (boiler), i is the annual interest rate,j is the equipment useful life (years) and t is the factory operationhours per year. Table 8 shows the data used to calculate the cost ofsteam cs
For all steam streams in the process (2.5 and 6 bar), as well asvapour and condensate streams, the samemonetary cost per unit ofexergy cs was adopted.
Fig. 4. Flow sheet of double-effect distillation process.
Table 4Main parameters for the simulation of conventional and double-effect distillation.
Parameter Conventionaldistillation
Double-effectdistillation
Top pressure, column D, (kPa) 134 25Top pressure, column BeB1, (kPa) 116 116Wine temperature at the inlet of
column A1, (�C)90 51.2
Condenser temperature of distillationa �
35 34
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Regarding the cost of cooling water (cold utility), the value re-ported by Ref. [16] was updated, resulting in ccw¼ 0.055 US$/m3
Thus the cost of utilities Cutility was calculated for each heaterand cooler of the production system, according to the followingequations:
Cutility ¼Xn
i¼1
Cheater;i þXm
j¼1
Ccooler;j (6)
Cheater;i ¼ Cs;i � Cc;i ¼ cs _ms;i�exs;i � exc;i
�(7)
Ccooler;j ¼ ccwvj _mcw;j (8)
Finally, the total annual cost is calculated as follows:
Fig. 5. Two-stage compression systemwith intercooling-compression station COMP-S.
Pequipment ¼ t _Ztotal ¼ tX
_Zi ¼ tEtotalx (9)
Putility ¼ tCutility (10)
set, ( C)Isentropic efficiency of pumps,b (%) e 70Isentropic efficiency of compressors,c
(%)e 80
Mechanical efficiency of compressors,c
(%)e 98
Intermediate pressure,d (kPa) e 57.2Outlet temperature of intercooling,
(�C)e 80
Washing column pressure, (kPa) e 101.3
a Condenser R.b B1, B2, B3 and B4.c COMP1, COMP2 and COMP-D.d Compression system of two stages COMP-S.
Table 5Streams adopted for thermal integration.
Hot streams Ti(�C)
Tf(�C)
DH*(MW)
Cold streams Ti(�C)
Tf (�C) DH*(MW)
Sterilized juice 130 32 41.8 Imbibition water 25.0 50 4.4Phlegmasse 103.8 35 3.0 Treated juice 34.2 105 44.1Vinasse 47.4 35 6.3 Pre-heating juice 98.1 115 2.6Anhydrous
ethanol78.3 35 8.9 Juice for
sterilization95.7 130 14.8
Vapourcondensates
112.4 35 8.8 Reboiler columnA
69.9 70 18.3
Condenser R 46.4 34 20.6 Reboiler column BeB1
103.8 103.81 23.5
Condenser columnEXTRACTa
78.3 78.3 7.6 Reboiler columnEXTRACTa
141 141 7.2
Reboiler columnRECUPERAb
149.6 149.6 2.5
a EXTRACT: extractive column in dehydration step.b RECUPERA: recovery column in dehydration step.
Table 8Data for determination of monetary steam cost.
Parameter Value
Boiler capital costa Eb (106 US$) 19.7Sugarcane production cost,b Ccane (US$/t) 29.4Pressure of steam generated in boiler, (bar) 90Temperature of steam generated in boiler, (�C) 520Boiler feed water temperature (�C) 116Boiler efficiency (%) 86Specific exergy of bagasse,c exbag (kJ/kg) 9885Specific exergy of sugarcane,c excane (kJ/kg) 5695Annual interest rate,b i (%) 15Equipment useful life,b j (years) 25Factory operation hours per yearc 4000
a Updated from Refs. [20] and [21].b Albarelli [21].c Ensinas et al. [22].
Table 6Main specifications in cogeneration system [6].
Cogeneration system
Pressure of boiler live steam, bar 100Temperature of boiler live steam, �C 530Isentropic efficiency of electricity generation steam turbines, % 80Alternator efficiency of turbine generator, % 97.6Turbine mechanical efficiency, % 98.2Pump isentropic efficiency, % 70Boiler thermal efficiency, % (LHV base) 86Mechanical power demand of cane preparation and extraction system,
kWh/t of cane16
Electric power demand of sugar and ethanol process, kWh/t of cane 12
Table 7Equipment cost in each sector of the factory for Base Case.
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Ptotal ¼ Pequipment þPutility (11)
3. Results and discussion
Table 9 presents the main results of simulation related todistillation process. It can be observed that in cases DED and DEDTIthe washing of gases with phlegmasse maintained the ethanollosses to the same level as the Base Case. This recovery is reflectedin a slight increase in anhydrous ethanol production at the expenseof second-grade alcohol. For this purpose, the amount of phleg-masse recirculated was 3800 kg/h (recirculation of 9.3%).
Regarding the energy consumption, in double-effect distillationcases (DED and DEDTI), the energy recovered from the integrationof the condenser of column BeB1with the reboiler of column Awas26.6 MW, hence the final energy consumption in the reboiler ofcolumn A was 19 MW. This represents an energy saving ofapproximately 58% only in column A.
Regarding the steam consumption in the distillation step (col-umn A and BeB1), the thermal integration of the overall process(DEDTI) promotes a further decrease, in comparison to the casewithout thermal integration (DED), according to Table 9.
Fig. 7a presents the Grand Composite Curve-GCC constructedfrom process streams, without including the MEE streams (Step 1).The horizontal bars (boxes) at the left of the graph represent the
Fig. 6. Scheme of cogeneration system: configurations I and II.
Table 9Main results of simulation.
Base DED DEDTI
Anhydrous ethanol (L/t of cane) 81.9 82.6 82.6Ethanol recovered in alcohol of second grade (L/t of cane) 1.5 0.79 0.79Ethanol loss in gases stream, (L/t of cane) 0.3 0.3 0.3Total losses in distillation, (%) 0.60 0.65 0.65Steam consumption in distillation, (kg/L of ethanol) 2.67 1.67 1.02Steam savings in distillation,a (%) e 37 62
a In comparison to the Base Case.
Fig. 8. Final GCC, including evaporation system streams.
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effects of theMEE system (each horizontal bar represents one effectof the MME). According to these graphs, whereas the standardoperating pressures in the evaporation system are adopted (Fig. 7a),the fifth effect intersects the GCC below the Pinch Point. For thisreason, the operation pressures of the evaporation system weremodified in order to fit the evaporation systemwithout intersectingthe GCC. This option allows the integration of the evaporationsystem without any additional consumption of hot utility [17,19].
In this study, the operating pressures of the 4th and 5th effectswere raised from 0.54 bar and 0.16 bar to 0.71 bar and 0.51 bar,respectively. Then, the temperature of the 5th effect resulted in81.3 �C, which is higher than the Pinch Point temperature (73.3 �C).
Fig. 7b shows the evaporation system with modified pressures.It can be observed that it fits above the Pinch Point and no effectintersects the GCC. According to this proposal, the vapour gener-ated in the 5th effect is used to cover heat demands in part of theprocess. In this way, the process serves as a condenser.
Fig. 8 shows the final GCC, including the streams of the MEEsystem with the modified pressures. Finally, targets of minimumenergy consumption can be determined from the final GCC.
Fig. 9 shows the heat exchanger network for maximum energyrecovery built in this work. Heat exchangers 10, 11 and 12 areheaters that use hot utility, while heat exchangers 13, 14 and 15 arecoolers that use cold utility. Streams presented in Table 5 that arenot presented in Fig. 9 have their heating or cooling demandcovered by only heat or cold utility. This heat exchanger networksatisfies the target of maximum energy recovery. However, it is notoptimized in terms of device cost yet.
Fig. 7. Grand composite curve (G
Table 10 shows the steam consumption of each operation of theproduction process, determined by the heat exchanger networkdesign (Fig. 9).
Table 11 shows the electricity and bagasse surplus for theevaluated cases. To calculate the electricity surplus, an electricconsumption of 16 kWh/t of cane was assumed in the juiceextraction system (including cane preparation devices), while forthe rest of the process 12 kWh/t of cane was assumed. The addi-tional electric consumption in double-effect distillation, owing tocompressors and pumps, was also estimated, and resulted in3.8 kWh/t of cane.
The results show that the electricity surplus in cases DED andDEDTI is significantly lower in comparison to the Base Case forConfiguration I, because of the reduction in steam consumption. Onthe other hand, the bagasse surplus increased significantly, which isinteresting if the ethanol production, based on second-generationtechnologies, is considered.
In Configuration II, when all the bagasse is burnt to producesteam, the lower steam consumption of cases DED and DEDTI
CC) þ evaporation system.
Fig. 9. Heat exchanger network built in this work.
Table 10Steam consumption of each operation of the process (kg/t of cane).
Base DED DEDTI
Steam at 6 barSterilization of must 51.1 51.1 7.4Dehydration: extractive column 24.8 25.0 24.9Dehydration: recovery column 8.5 8.5 8.6Steam at 2.5 barJuice treatment e e 32.6Pre-heating juice e e 8.6Evaporation system 164.3 164.3 46.9Distillation column A 147.1 60.5 6.8Distillation column BeB1 71.7 77.4 77.6
Total 467.5 386.9 213.4Savinga (%) e 17 54.3
a Reduction in steam consumption in comparison to the Base Case.
Table 12Monetary cost per unit of exergy.
Monetary cost per unit of exergy (10�6 US$/kJ)
Sugarcane 5.16Steam 18.6
R. Palacios-Bereche et al. / Energy 82 (2015) 512e523 521
promotes greater electricity surplus because a greater amount ofsteam passes through the electricity generation turbines.
The results obtained in this study can be compared with othersfrom the literature. For instance, Junqueira et al. [1] reported a 36%reduction in steam consumption (in overall process), which ishigher than the reduction obtained in Case DDE (17%), since [1]assumed the distillation triple effect in combination with vacuumextractive fermentation. On the other hand, it is interesting to notethat Case DDETI presents a steam consumption lower than thatpresented by Ref. [1]. Bessa et al. [5] reported a reduction in steamconsumption of 54% only in the distillation step, while in the
BPST: Back pressure steam turbines; CEST: jCondensing extracting steam turbines.a Percentage in relation to the total bagasse produced.
present work, this reductionwas 37%. It can be explained by the factthat [5] assumed a total condenser in a high pressure column,where hydrated ethanol is obtained in a liquid phase. In contrast, inthe present study, a partial condenser was assumed since hydratedethanol in the vapour phase is fed to the dehydration column. Forthis reason, in the present study, there is low energy available forrecovery.
Regarding the economic assessment, the diameters of columnsD, A and A1 in the double-effect distillation system (cases DED andDEDTI) were approximately 50% higher than the respective di-ameters in the Base Case. This led to a cost increase of 28.6%, withreference to these columns. Moreover, the costs of additional de-vices, such as the compressors and the absorption column, weretaken into account. Thus the final cost in the double-effect distil-lation system resulted in 2.69 times the cost of a conventionaldistillation system.
R. Palacios-Bereche et al. / Energy 82 (2015) 512e523522
For the thermal integrated case DEDTI the cost of a new heatexchanger network (Fig. 9) was 2.3 times that of the Base Case.Therefore, the total equipment cost estimated for cases DEDand DEDTI was 153.3 � 106 US$ and 157.4 � 106 US$,respectively.
For the cost of utilities, Table 12 shows the monetary cost perunit of exergy for sugarcane and steam calculated according to theprocedure in Section 2.6.
Fig. 10 shows the annual costs with reference to the equipmentand the utilities. Cases DED and DEDTI showed a significantreduction in the cost of utilities (16% and 52%, respectively, incomparison to the Base Case), but on the other hand, they showed aslight increase in investment cost (7% and 10%, respectively). Itmust also be mentioned that the equipment cost is significantlyhigher than the cost of utilities. However, the sum of these costs(equipmentþ utilities), in case DEDTI, was 11% lower in comparisonto the Base Case, while in case DED, there was no significantdifference.
The simulation of the ethanol production process, includingdouble-effect distillation technology, was performed in this study.The incorporation of double-effect distillation in the productionprocess promoted a significant reduction in the steam consump-tion of the process. Moreover, the thermal integration procedurethrough the Pinch Method was applied to the whole process, inorder to achieve a further reduction in steam consumption of theprocess (DEDTI case). Thus, cases DED and DEDTI presented areduction in steam consumption of the overall process, of 17% and54%, respectively, in comparison with the Base Case. The thermalintegration procedure also allowed the determination of the mostappropriate operation pressures in the MEE system (appropriateplacement), to minimize the energy consumption of the process.Regarding the cogeneration system, the decrease in steam con-sumption led to a reduction of 33% and 81% in the electricitysurplus, in comparison with the Base Case, when back-pressuresteam turbines were adopted. In contrast, electricity surplus in-creases of 3% and 22% can be achieved when extrac-tionecondensing steam turbines are adopted. For the economicassessment, the additional investment cost, owing to the imple-mentation of a double-effect distillation system and thermalintegration, was estimated. The results of the economic assess-ment showed that DEDTI Case is the most advantageous becausethe equipment cost together with the cost of utilities was 11%lower in comparison to the Base Case, while in the DED case therewas no significant difference.
Acknowledgements
The authors wish to thank CNPq (Process PQ 304820/2009-1)for the researcher fellowship and the Research Project Grant (Pro-cess 470481/2012-9), and FAPESP for the Post-Ph.D. fellowship(Process 2011/05718-1) and the Research Project Grant (Process2011/51902-9).
Appendix
Table A1 shows data for the main stream data in the distillationprocess.
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