Technical Performance and Economic Evaluation of Evaporative and

Sep 18, 2017 - This study evaluated the trade-offs in product recovery, energy consumption, and economics between evaporative and membrane-based ...
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Technical Performance and Economic Evaluation of Evaporative and Membrane-Based Concentration for Biomass-Derived Sugars David A. Sievers,* Jonathan J. Stickel, Nicholas J. Grundl, and Ling Tao National Bioenergy Center, National Renewable Energy Laboratory, Golden, Colorado 80401, United States S Supporting Information *

ABSTRACT: Several conversion pathways of lignocellulosic biomass to advanced biofuels require or benefit from using concentrated sugar syrups of 600 g/L or greater. While concentration may seem straightforward, thermal sugar degradation and energy efficiency remain major concerns. This study evaluated the trade-offs in product recovery, energy consumption, and economics between evaporative and membrane-based concentration methods. The degradation kinetics of xylose and glucose were characterized and applied to an evaporator process simulation. Although significant sugar loss was predicted for certain scenarios due to the Maillard reaction, industrially common falling-film plate evaporators offer short residence times (100 g/L. A second step using evaporation is necessary to achieve target concentrations. Techno-economic process model simulations evaluated the overall economics of concentrating a 35 g/L sugar stream to 600 g/L in a full-scale biorefinery. A two-step approach of preconcentrating using membranes and finishing with an evaporator consumed less energy than evaporation alone but was more expensive because of high capital expenses of the membrane units.

1. INTRODUCTION Sugars obtained from lignocellulosic biomass are an attractive feedstock for renewable fuels, chemicals, and materials.1,2 Many pathways for the upgrading of cellulosic sugars to advanced biofuels require (or benefit from) concentrated sugar solutions (syrups).3,4 Previous work identified and developed an economical solid−liquid separation process for slurries obtained from batch enzymatic hydrolysis (BEH).5 Flocculation and vacuum filtration remove the solids, and the clarified sugar stream (“hydrolysate”) has ∼75 g/L total sugars after sufficient washing of the cake to recover 95% of the sugars. A novel continuous enzymatic hydrolysis (CEH) process has also been proposed that integrates solids removal by membrane filtration.6 The clarified hydrolysates from CEH are 20−40 g/L biomass-derived sugars, depending on the process design. However, downstream processing steps via fed-batch fermentation and inorganic catalysis significantly benefit from consuming concentrated sugar syrups rather than native concentration hydrolysis liquors, for which specific productivity is increased and subsequent equipment size may be reduced. The preferred concentration of the sugar is typically 400−700 g/L depending on the upgrading pathway. Therefore, the hydrolysate must be concentrated prior to further conversion. Evaporation is a well-characterized industrial process commonly used to remove water and concentrate solutions of various types, including sugar solutions in the food industry.7 Evaporation is simple conceptually, and several commercial© XXXX American Chemical Society

scale equipment designs are available to suit specific applications. As water is vaporized from the product, so are volatile organic compounds produced during upstream diluteacid thermochemical hydrolysis (pretreatment). Some of these compounds include bioconversion inhibitors, such as furfural, hydroxymethylfurfural (HMF), and acetic acid, and their removal is a benefit. The primary disadvantage of evaporation is the enormous energy required to induce the phase change of water from liquid to gas, either by raising the temperature and/ or by pulling a vacuum. Additionally, the elevated processing temperatures also facilitate sugar degradation reactions, possibly to the detriment of product quality and overall yield. Some researchers have demonstrated that acid hydrolysis, caramelization, and Maillard reactions are possible, depending on the specific sugars involved and the evaporation conditions.8−10 Although others have characterized the degradation kinetics of several sugars including glucose,11−15 these studies were all on pure sugar solutions, not real biomass hydrolysates that contain a complex mixture of sugars, proteins, and other uncharacterized soluble components. Furthermore, most of the previous studies were performed at temperatures elevated far above the relevant range for evaporation. Received: Revised: Accepted: Published: A

May 26, 2017 September 12, 2017 September 18, 2017 September 18, 2017 DOI: 10.1021/acs.iecr.7b02178 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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The prepared liquors were refluxed in a 250 mL boiling flask, heated with an electric mantle set at ∼65 W power, and agitated using a magnetic stir bar on a stir plate at 350 rpm. Vacuum was applied during some experimental runs at 33 or 50 kPa, and some runs were executed at atmospheric pressure (∼81 kPa absolute pressure at an altitude of 1800 m), resulting in an overall boiling point range of 70−100 °C. The presence of sugars in solution elevated the boiling points of solutions from that of pure water. Liquors were refluxed for 6−24 h (or sometimes longer for low starting concentrations), and timecourse samples were analyzed for pH, for sugar and furan concentration by high-performance liquid chromatography (HPLC),22 and for discoloration on a NanoDrop 8000 spectrophotometer (Thermo-Fisher, Wilmington, DE). Kim and Lee23 demonstrated a linear relationship between Maillardassociated melanoidin formation and absorbance at 420 nm. Rather than assemble complicated continuous concentration equipment on the bench and determine absolute sugar degradation yield, sugar degradation in scaled-up equipment was simulated using a kinetics rate model and experimentally obtained rate constants. The Maillard reaction is quite complex and involves multiple steps between sugars, peptides, intermediate products, and even water. Although simplification to a classic zero-, first-, or second-order reaction does not comprehensively describe the kinetics, it is useful as an engineering tool, and results can be used to predict the system within the range of experimental conditions.24 The degradation of sugar was modeled using first-order rate kinetics:

An alternative to evaporation for solute concentration is membrane separation. Reverse osmosis (RO) or nanofiltration (NF) membranes can be used to retain the small molecules of interest (sugars) while permeating (removing) water. Advances in membrane technology have made membrane concentration an attractive means of concentrating fruit and vegetable juices in recent years. 16,17 Membrane concentration may be performed at relatively low temperature and does not impose a phase change; hence, the energy costs are lower. Malmali et al.18 recently examined the concentration of hydrolysate sugars by NF membranes, demonstrating high retention of sugars and reasonable permeance. In addition, they demonstrated clearance (removal) of some inhibitory organic compounds, including acetic acid, furfural, and HMF. However, membrane permeance decreases at higher sugar concentrations because of increased viscosity. More recently, membrane distillation has been proposed for concentrating sugar solutions.19 A membrane is used to facilitate the phase change in membrane distillation, lowering the energy cost requirements. Because membrane distillation is still in early stages of technology development, it was not evaluated here. In this work, the concentration of biomass-derived sugars using evaporation and membranes was characterized experimentally, and the results were evaluated for industrial use. Xylose and glucose degradation and loss were studied, and the reaction kinetics were characterized. Sugar preservation yields of a biomass-derived sugar solution were simulated within relevant industrial evaporator equipment using experimental sugar degradation kinetics data. Membrane concentration performance was also characterized using the same sugar solution. These data were used in a techno-economic analysis (TEA) model to compare industrial evaporation and membrane concentration routes within a biorefinery in order to understand which is more economically viable.

dci = −kici dt

(1)

where ci is the sugar concentration (i = x for xylose, i = g for glucose, g/L), t is time (min), and ki is the condition-dependent rate coefficient (min−1) that is affected by temperature, pH, and possibly other factors. The kinetics of both xylose and glucose degradation were assessed by heat-treating solutions and observing their behavior. First-order rate coefficients were derived from the reflux experiment results, and xylose and glucose loss rates were calculated for specific conditions of interest using eq 1. Calculated loss rates were used in timeevolution reaction simulations to estimate the amount of sugar degradation and resulting recovery yield given equipmentspecific residence times. 2.2. Membrane Concentration. A high-pressure membrane-testing unit (HP4750, Sterlitech Corporation, Kent, WA) was used to test solute rejection and permeance of various membrane materials. The HP4750 is designed to use membrane disks that are 49 mm in diameter, with an effective surface area of 14.6 cm2. The HP4750 unit was placed in a water bath on a heated stir plate to control temperature. Pressure was applied from a high-pressure nitrogen-gas cylinder equipped with a pressure regulator. The permeate was collected in an Erlenmeyer flask placed on a mass balance. The mass reading was recorded at defined intervals on a computer using serial communication. Permeate flow rates were subsequently calculated by taking numerical derivatives of the mass and time signals. Temperature was measured with a thermocouple and recorded at defined intervals with a QuadTemp2000 data logger (MadgeTech, Warner, NH). Malmali et al.18 identified several promising membranes for concentrating sugar solutions. Three membranes were retested here: NF270 (Dow Filmtec, Edina, MN) and RO99 and RO90 (Alpha Laval, Lund, Sweden). The dry, as-received, membranes

2. EXPERIMENTAL METHODS 2.1. Sugar Degradation Kinetics. Corn stover was deacetylated using 0.4 wt % sodium hydroxide and soaked in 1.0 wt % sulfuric acid as described by Chen et al.;20 it was then thermochemically hydrolyzed at 160 °C for 15 min. This material was subsequently enzymatically hydrolyzed at 50 °C for 3 days using 20 mg of Novozymes CTec2 per kilogram of cellulose starting at 20% total solids, yielding 41 and 65 g/L xylose and glucose, respectively. This slurry was centrifuged and sterile-filtered, and 5 g/L glycerol was added as a chemically inert tracer. Liquor was used as-is or concentrated 5-fold in a vacuum rotovap at 30 kPa absolute pressure and resulting boiling temperature (∼70 °C), which subjected the sugars to heat for about 1 h with no detectable degradation. Protein content of the native liquor was quantified by using elemental nitrogen as a proxy.21 Dried liquor solids contained 0.15% nitrogen, translating to a protein concentration of approximately 6.9 g/L. Native pH was 4.8, and some of the liquors were adjusted to pH ∼3 with 10% sulfuric acid prior to use. Concentrated synthetic sugar liquor was prepared with 200 g/L xylose and 330 g/L D-glucose in a 110 mM citrate buffer. pH was adjusted to ∼3.8 using 10% NaOH. In some cases, 50 g/L Difco bacto casamino acids (a hydrolyzed casein product with 10% nitrogen) was added to simulate the protein nitrogen concentration found in real hydrolyzate. Again, glycerol was dosed at 22 g/L as a chemically inert tracer. B

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Figure 1. Process flow diagrams for production of hydrocarbon fuel from lignocellulosic feedstock. Sugar intermediates are concentrated into syrups by evaporators for downstream use (a), or with the addition of a membrane concentrator step to reduce size and energy demand of the subsequent evaporators (b). Evaporation and membrane concentration experimental results were used to model the behavior and economics of these processes. Waste-water treatment and utilities not depicted.

rejection vary with retentate concentration, i.e., ψ = ψ(c) and r = r(c). The membranes experience compaction during the conditioning; hence, they exhibit a decrease in permeance during the time-course of conditioning and approach a steady-state permeance value. The measured water permeance of the three membranes during conditioning is shown in Figure S1 in the Supporting Information. Because of the closed experimental dynamic-batch system, it is impractical to take run-time measurements of the retentate sugar concentration. This poses a challenge for determining how permeance and membrane rejection relate to retentate concentration. However, the measured experimental values of permeate concentration and collected permeate mass could be used, along with an empirical rejection versus concentration model, to determine the run-time retentate concentrations. The rejection model is given by

were initially soaked in deionized (DI) water for at least 1 h. They were subsequently conditioned by flowing DI water through them at an applied pressure of 30 bar and temperature of 40 °C. Membranes were stored in DI water in a refrigerator when not in use. The performance of the membranes were tested with either a synthetic glucose solution (50 g/L glucose, 25 mM sodium citrate, pH 5.0, electrical conductivity [EC] 6.6 mS/cm) or with clarified and diluted hydrolysate liquor from deacetylated, dilute-acid, and enzymatically hydrolyzed corn stover, as described in section 2.1 (44.4 g/L total sugar, 26.1 g/L glucose, 15.1 g/L xylose, pH 4.6, EC 1.5 mS/cm). Applied pressure was 30 bar and temperature was 40 °C for all experiments. Run-time samples of permeate were collected periodically during the concentration experiments. Samples of pooled end point permeate and retentate were also collected. All samples were analyzed for typical hydrolysate sugars by HPLC using NREL standard procedures.22 The performance of membranes for concentrating solutes may be characterized by permeance (L/(m2·h·bar) or “LMH/ bar”), defined as ψ=

J ΔP

r = a0 − e−a1c a2 = a0 − exp(−a1 + a 2 ln c)

(4)

where ai are empirical coefficients; the numerical methods for computing the retentate concentration are described in the Supporting Information. An empirical model for permeance

(2)

and rejection (unitless) of solutes (sugars here), defined as cp r=1− (3) c

ψ = b3 + 0.5b2 erfc[b1(c − b0)]

(5)

was also fit to the data, where bi are empirical coefficients. As these models are empirical and only intended to reproduce the experimental trend lines, their form and coefficient values should not be used to infer any physical significance. The models for rejection and permeance were subsequently used in the techno-economic models to calculate needed membrane area, flow rates, retentate concentrations, etc.

In these equations, J = Qp/A is permeate flux (L/(m ·h) or “LMH”), ΔP the pressure drop across the membrane (simply applied gauge pressure here in bar), Qp the permeate flow rate (L), A the cross-sectional area of the membrane (m2), cp the run-time (instantaneous) permeate concentration (g/L), and c the retentate concentration (g/L). Both permeance and 2

C

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Industrial & Engineering Chemistry Research Table 1. Reflux Experimental Conditions and Resulting Degradation Rate Coefficients liquor source synthetic synthetic biomass biomassa biomassa biomassa biomass biomass biomass biomass biomass biomass biomass a

peptide content none added native native native native native native native native native native native

absolute pressure (kPa) 81 81 81 81 81 81 30 81 81 47 47 30 30

temperature (°C) 98.7 99.0 95.7 95.3 96.8 96.7 70.0 100.0 99.5 83.4 85.3 73.5 73.5

starting pH

starting total concentration (g/L)

3.78 3.85 2.91 4.78 4.79 4.78 4.82 3.17 4.78 2.93 4.87 2.72 5.04

521 520 100 101 102 102 101 533 549 562 546 553 560

t (h) 24 6 24 24 24 24 24 6 6 6 6 6 6

kx (min−1) 5.87 4.84 1.46 5.66 5.57 5.27 2.40 7.15 1.06 6.81 2.25 3.02 1.65

× × × × × × × × × × × × ×

−5

10 10−4 10−6 10−5 10−5 10−5 10−5 10−4 10−3 10−5 10−4 10−5 10−4

kg (min−1) 2.05 1.59 5.65 1.53 1.90 1.40 2.35 2.36 3.20 8.39 7.83 2.04 6.91

× × × × × × × × × × × × ×

10−5 10−4 10−6 10−5 10−5 10−5 10−5 10−4 10−4 10−13 10−5 10−6 10−5

Replicated condition.

Figure 2. Concentration (a) and absorbance (b) of pure sugar solutions over reflux time.

reactors (CSTR) with integrated product filtration produced a clarified cellulosic sugar stream. Four CEH reactor units were placed in series, all operating with 7.5% insoluble solids and enzyme dosing of 10 mg protein per g cellulose into the first reactor. CEH achieves high productivity but results in a relatively dilute clarified sugar stream on the order of 34 g/L total sugar, because the soluble sugar product is continuously diafiltered out of the system using water. The dilute sugar stream is concentrated to 637 g/L using three vacuum evaporation stages (18× concentration factor in total). This process layout is called the base case and is shown in Figure 1a. The sugar loss rate during evaporation at elevated temperature was calculated based on experimental results and incorporated into the simulations. An alternative case was also evaluated using a membrane separator unit inserted prior to the evaporators (shown in Figure 1b) to reduce the required duty of the evaporators and was evaluated at multiple levels of implementation. In this alternative case, various levels of preconcentration were performed by the membrane unit (ranging between 50 and 100 g/L intermediate concentration) prior to final concentration by evaporators. The membrane permeate volumetric flow and sugar rejection were estimated based on experimental results, membrane equipment specifications, and retentate sugar product concentration target. Required membrane area was calculated using mass flows to achieve the required outlet concentration and the experimentally determined permeance

3. TECHNO-ECONOMIC ANALYSIS METHOD To conduct a fair comparison between concentration technologies using the experimental data, a TEA process model was built based on NREL’s recent model for bioconversion of lignocellulosic biomass to hydrocarbon liquid fuels at a scale of 2000 Tonne/day dry feedstock. This model described by Biddy et al.1 is the “fuels-only/improved case” (no coproducts and using updated design details), which deacetylates the feedstock followed by dilute-acid thermochemical hydrolysis (pretreatment), batch enzymatic hydrolysis, and filtration of the produced sugars prior to concentration using a falling-film plate evaporator heated by mechanical vapor recompression (MVR). Fed-batch aerobic bioconversion by oleaginous yeast using the sugar syrup produces intracellular lipids, which are extracted and upgraded to renewable diesel blendstock and naphtha fuels. The process was simulated using Aspen Plus to determine mass and energy balances for each unit operation of the entire layout. Operating expenses were combined with the total capital investment in a discounted cash flow rate of return analysis to determine a minimum fuel selling price (MFSP), which was the minimum price that yielded a zero net present value for a 10% rate of return. The version of this process layout used for our analysis here replaces the sequential batch enzymatic hydrolysis step with a continuous approach to enzymatic hydrolysis (CEH), which combines the filtration step within this unit operation. In the CEH step, multiple sequential continuously stirred hydrolysis D

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Figure 3. Concentration (a) and absorbance (b) of biomass-derived sugar solutions over reflux time. The averages of the replicated condition were plotted (97 °C, pH 4.8, 100 g/L).

Lu13 studied the acid-catalyzed hydrolysis of pure xylose between 180−220 °C, and when their Arrhenius-based kinetics parameters are applied to a temperature of 100 °C as in our work, the hydrolysis reaction rate is negligible. Caramelization reactions and associated browning can take place in the absence of peptides and are also usually associated with Maillard reactions. Although slight browning of the pure sugar did occur after 24 h, no color change was measured after only 6 h, leading to the conclusion that the Maillard reaction (when peptides are present) is of greater importance in this system than caramelization. Additionally, Lan et al.9 noted caramelization was prevalent in a system containing fructose but could be neglected in favor of the Maillard reaction in some systems involving only xylose and glucose. Moving to real biomass-derived liquors, the effects of temperature and starting pH were studied, and more evidence to support the predominance of the Maillard reaction was garnered. Figure 3 plots the sugar concentrations and browning of the solutions versus time. Furfural and HMF generally increased in concentration, and absorbance increased as sugars were lost. Relatively low degradation rates were observed on the native-concentration liquor, even with varying pH. The first-order rate coefficients of xylose and glucose loss were determined as 5.50 ± 0.23 × 10−5 and 1.61 ± 0.29 × 10−5 min −1 , respectively, for the replicated condition. The degradation rates of concentrated liquors were substantially higher, and the effects of temperature and starting pH were also observed. Figure 4 is a first-order polynomial response-surface model that demonstrates the effects of starting pH and reflux temperature on the xylose-loss rate coefficient for highconcentration reflux experiments. The rate coefficient for glucose loss was approximately 3.2× lower than that for xylose for all conditions. As expected, refluxed liquors under vacuum and at lower temperature generally resulted in lower sugar losses. Contradictory to expectations, the samples that were adjusted to a lower pH near 3 prior to reflux actually conserved sugars better than those of native pH around 4.8. This additional observation supports the theory that the Maillard reaction is the dominant mechanism of sugar degradation here, as lower pH has been shown to retard the Maillard reaction.28 The pH of refluxing liquors decreased over time, which is another trait of the Maillard reaction.

parameter, which sets the sizing basis for the unit. The sugar rejection data was used to adjust the mass fractions of sugars in the permeate/retentate to replicate experimentally observed values. In addition to membrane sizing, power requirements for the high-pressure feed pump (20−40 bar) were calculated. The operating and capital costs for both evaporators and membrane units were estimated using data from previous studies.25,26 Periodic membrane material replacement costs are also included in the operating costs, which occur on average every two years.

4. RESULTS AND DISCUSSION 4.1. Evaporative Concentration. A total of 13 reflux experiments were executed, with two of them using synthetic liquor and the remainder using real biomass-derived hydrolysate liquors. Table 1 further elaborates on the conditions of each reflux experiment. Three replicates at atmospheric pressure were performed with liquor at the native pH and concentration in order to estimate experimental uncertainties. Glycerol concentration remained constant for each experiment, indicating negligible evaporation of water. The synthetic sugar liquor without protein yielded a relatively low degradation rate with approximately 92% xylose and 97% glucose remaining after 24 h; more than 98% of each sugar remained and there was no detectable coloration at 6 h. After 24 h, furfural and HMF were present at 2.5 and 0.7 g/L, respectively. The same experiment with added peptides yielded a much higher rate of sugar degradation. Figure 2 demonstrates the stark difference between the liquors when there is a small amount of peptides available for reaction. After just 6 h with peptide present, xylose and glucose concentration dropped to 85% and 95%, respectively, and the slightly yellow solution progressed to a dark brown as indicated by the large increase in absorbance at 420 nm. Furfural and HMF were also measured at 7.4 and 2.0 g/L, respectively, after 6 h. This experiment revealed that the presence of protein (hence, amino acids) in solution at this concentration is an important factor affecting the loss of sugars. Although furans were produced, their presence does not directly suggest acidhydrolysis as a mechanism, as furfural and HMF are intermediate products of the common Maillard reaction between peptides and sugars.27 In fact, the progression of the Maillard reaction results in the formation of melanoidins and the browning of the solution,9,23 as observed here. Jing and E

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water vaporization. Figure 5 demonstrates the results for three scenarios. The batch evaporator subjects the sugars to a very long residence time of ∼30 h, and even at a reduced temperature of 70 °C, significant degradation occurs because of the requisite exposure time. Only 88% of the total sugars remain after concentration. The FCE was simulated at 100 °C (atmospheric pressure), and although the mean residence time of recirculation was only 2 h, significant degradation is still predicted. Fortunately, for the very short residence time needed to achieve the desired concentration in industrial-scale evaporation equipment, the predicted sugar degradation is negligible. Falling-film plate evaporators are heavily utilized industrially, and they both share a common theme of operation where preheated lean liquor is pumped to the top of a heat exchange surface and allowed to fall down while steam is applied on the other side, driving off vapor and producing a concentrated product at the bottom. These units operate in plug-flow fashion, and residence time within the heated section is on the order of 5 min.30 This short residence time appears to be the key to successfully concentrating biomass-derived sugars while preserving the sugar and quality of product, and the subsequent TEA here assumes no sugar degradation takes place. 4.2. Membrane Concentration. Nanoporous membranes offer an alternative separation mechanism for removing water and concentrating sugar solutions, with potential economic savings for cases where the initial sugar concentrations are low and large amounts of water must be removed. The technical performances of a few candidate membranes were evaluated for use in the techno-economic analysis. Summary results for membrane concentration experiments are shown in Figure 6. The RO90 and RO99 membranes exhibited similar permeance and sugar rejection, with the RO90 having slightly higher permeance when tested with the glucose solution. Although the RO99 membrane was not tested with hydrolysate, we expect that it will have proportionally lower permeance similar to that observed with the RO90 membrane. The permeances of the RO90 and RO99 membranes are quite low, resulting in a flux of less than 10 LMH with 30 bar of applied pressure when the retentate sugar concentration is ∼100 g/L. Electrical conductivity (EC), a proxy for total concentration of inorganic salts, was measured for the end

Figure 4. Response-surface map of xylose degradation rate coefficient versus temperature and starting pH for high-concentration reflux experiments (∼550 g/L total starting sugars). Glucose degradation coefficients were about 3.2× lower. The points are the experimentally measured values.

The Maillard reaction is very complex and still not completely understood.29 To simplify the analysis, a firstorder kinetics model was used to define sugar-loss rate, which resulted in satisfactory fitting with experimental data of each individual run. The rate coefficients were not consistent between differing sugar concentrations, even though the firstorder kinetics model (eq 1) includes a concentration term. Starting with a total sugar concentration of ∼550 g/L and native pH, rate coefficients were approximately 10−20 times higher than at the native concentration of 100 g/L. In order to model the degradation of sugars over time in a real evaporator system, the rate coefficients were interpolated between measured values to obtain a rate coefficient that varies with concentration. These interpolated rate coefficients were used at native pH (4.8) to model the expected final total sugar concentrations and corresponding yields for three scenarios: (a) a batch evaporator at 30 kPa absolute pressure, (b) a continuous forced-circulation evaporator (FCE) at atmospheric pressure, and (c) for the scaled-up case utilizing a large continuous industrial falling-film plate evaporator unit assumed to require 5 min of residence time to reach the desired final concentration.30 The concentration of a sugar solution was simulated for a 5-fold reduction in total volume of liquor by

Figure 5. Simulation results for total xylose and glucose concentration (blue) and overall yield (red) when concentrating 120 g/L solutions 5-fold. Results for a batch evaporator at 70 °C (a), continuous forced-circulation evaporator (FCE) at 100 °C (b), and a continuous industrial-scale fallingfilm plate evaporator at 70 °C (c) are shown. F

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Figure 6. Summary results: (a) permeance and (b) total sugar rejection vs retentate concentration. The symbols and solid lines are measured data, and the dashed lines are model fits.

Figure 7. Capital-cost (a) and annual operating-cost contributions (b) of sugar concentration for varying membrane outlet targets.

point retentate and pooled permeate in all experiments. In all cases, the EC of the permeate was very low, while the retentate EC was increased over the feed in proportion to the amount of concentration. This indicates that salts are not passing through these NF membranes. The reduction in permeance as sugar concentration increases is due to increased osmotic pressure from all dissolved species, including sugars and salts. While the RO270 membrane had a permeance that was roughly twice that of the RO90 membrane, it had significantly lower rejection of sugars. Interestingly, rejection was higher for the hydrolysate than for the simple glucose solution, which was also observed by Malmali et al.18 The best-fit empirical coefficients for the rejection and permeance models for each membrane and solution combination are given in Table S1. 4.3. Techno-Economic Analysis. Experimental data on the RO99 membrane was utilized in the alternative case process model (incorporating membrane concentration followed by evaporators, Figure 1b) to understand the cost and energy impacts of implementing membrane units for sugar concentration. Various retentate sugar concentrations (feed concentrations to evaporators) were studied in the TEA to explore the full range of experimentally tested conditions. In each simulation, the retentate sugar concentration of the membrane unit was set at a target value (50−100 g/L) with the remaining concentration to 637 g/L carried out by the evaporator. The capital and operating costs for both the membrane and evaporator units were analyzed, and integrated process MFSP values were calculated.

Given the current estimated performance and cost, use of a membrane concentrator upstream to reduce the size and energy use of the subsequent evaporator does not provide economic improvements in the alternative case (Figure 1b) over sole use of the evaporator in the base case (Figure 1a). Compared to the base case, calculated to result in $10.03/GGE (2011 USD per gasoline-gallon-equivalent amount of energy), implementation of the RO99 membrane at the lower retentate concentration duty of 50 g/L yielded a 4% increase in MFSP at $10.41/GGE. When operated at the upper technical limit to produce 100 g/L sugar for evaporator feed, the MFSP increased by 42% over the base case to $14.23/GGE. The operating and capital expense costs for sugar concentration by combined membrane separation and evaporation approaches are shown in Figure 7. The increases in MFSP with higher evaporator feed sugar concentration (greater preconcentration by membranes) are predominantly due to membrane equipment capital costs, with the more favorable operating costs initially dampening the MFSP increase, but without an overall advantage. As seen in Figure 7b, operating cost for the sugar concentration step actually decreases when operating at low retentate concentrations (50−70 g/L), mainly due to relatively small required area for the membrane and coupled with lower membrane material replacement costs. The cost reduction is a result of the reduced power consumption of the combined membrane−evaporator system relative to the base case (evaporator only). The 2000 Tonne/day base case models evaporator electrical consumption at 26.9 MW, where if G

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sugar before final evaporation. Overall costs were elevated using the two-stage approach because of membrane capital and operating expenses. Continuous efforts on improving both membrane performance and cost are necessary to make membranes a cost-effective option for biomass-derived sugar concentration; particularly effective would be achieving higher permeance values. While not economically practical, membranes do improve the energy and sustainability footprint by reducing energy usage of the MVR-heated falling-film plate evaporator. Therefore, this could translate to significant life-cycle analysis gains with either more green electricity exported to the grid or with less demand for grid electricity.

the sugar stream is preconcentrated using membranes from 34 to 50 and 100 g/L, total electrical consumption between the membrane concentrator and smaller evaporator should decrease to 18.4 and 8.4 MW, respectively. As the membrane outlet setting (evaporator feed) rises above 80 g/L, the total operating cost starts to increase because of increasing membrane replacement costs and diminishing power savings. As the retentate concentration increases, the grid electricity demand shifts from a net import cost of $4.59 MM/yr (2011 USD) in the base case to a net export profit ranging from $0.25−5.83 MM/yr (grid purchase and selling prices assumed to be $0.07/kWh and $0.06/kWh, respectively.) Figure 7a shows the capital cost of evaporator, membrane, and sum of both with variation of the membrane retentate (product) sugar concentration setting. The evaporator capital decreases slowly as the retentate concentration setting is increased, whereas the membrane concentrator capital increases at a rapid rate. The evaporator capital decreases only slightly because the equipment is sized and costed by the flow rate going into the evaporator. With higher retentate concentration, the total flow decreases. The membrane capital increase is due to the nonlinear permeance response and associated membrane area requirements. The combined capital costs of evaporator and membrane is higher than the capital cost of the referenced baseline for all cases. A membrane concentrator upstream of evaporation offers reduction in the total process energy requirements, offsetting significant portions of the MVR-heated evaporator electrical demand. However, the large capital expense of the initial membrane system installation and membrane maintenance (replacement) operational costs are projected to eliminate the economic advantages unless these cost can be reduced by the manufacturers. As discussed earlier, osmotic pressure is a limitation of the tested membranes, and the resulting low permeance drives up the required size (area) and cost of the membrane concentrator unit. To achieve cost parity at the 100 g/L case with the base case, a 300% improvement in permeance would be required, increasing from 0.24 to 0.98 LMH/bar. This technical performance improvement would allow cost-competitive use of membranes at current prices to reduce the overall energy demand for this sugar concentration process.

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ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.7b02178. Detailed analysis of membrane-concentration data for the experimental work (PDF)

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Phone: +1 (303) 384 7748. Fax: +1 (303) 384 6877. ORCID

David A. Sievers: 0000-0002-7471-460X Ling Tao: 0000-0003-1063-1984 Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This work was funded by the U.S. Department of Energy through the Bioenergy Technologies Office under Contract No. DE-AC36-08-GO28308 with the National Renewable Energy Laboratory. The publisher, by accepting the article for publication, acknowledges that the U.S. Government retains a nonexclusive, paid up, irrevocable, worldwide license to publish or reproduce the published form of this work, or allow others to do so, for U.S. Government purposes.

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5. CONCLUSIONS The protein-catalyzed Maillard reaction was identified as the primary route of sugar degredation in hydrolysates when exposed to temperatures required for evaporative concentration. Although significant sugar loss was predicted in certain scenarios, concentrating sugars using an industrially relevant method that minimizes residence time at evaporation temperatures alleviated concerns for degradation and yield loss. The preferred method of concentration is the falling-film plate evaporator, which inherently has a short residence time leading to negligible predicted sugar losses. Membrane concentration experiments revealed that sugars exhibited very high rejection (near unity) for some of the membranes; however, permeance decreased significantly with concentration and was relatively low when the retentate sugar concentration reached 100 g/L (and higher). To achieve the target of at least 600 g/L sugar syrup, membrane concentration would also require a second step, such as an evaporator, to complete the process. Although overall energy consumption was lower than using only evaporation, TEA revealed no economic benefit from using membranes to preconcentrate

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DOI: 10.1021/acs.iecr.7b02178 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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DOI: 10.1021/acs.iecr.7b02178 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX