Scholarly article on topic 'Using pyrolytic acid leaching as a pretreatment step in a biomass fast pyrolysis plant: Process design and economic evaluation'

Using pyrolytic acid leaching as a pretreatment step in a biomass fast pyrolysis plant: Process design and economic evaluation Academic research paper on "Chemical engineering"

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{"Fast pyrolysis" / "Acid leaching" / "Alkali and alkaline earth metals" / "Pyrolytic sugars" / "Process design" / "Techno-economic evaluation"}

Abstract of research paper on Chemical engineering, author of scientific article — S.R.G. Oudenhoven, A.G.J. van der Ham, H. van den Berg, R.J.M. Westerhof, S.R.A. Kersten

Abstract Removing alkali and alkaline earth metals (AAEMs) from biomass, with pyrolytic acids, before pyrolysis leads to increased organic oil and sugar yields. These pyrolytic acids are produced and concentrated within the pyrolysis process itself. The purpose of this paper was to evaluate under which conditions acid leaching of pinewood, bagasse and straw can improve the technical and economic feasibility of a pyrolysis process. Therefore, a preliminary process design for the implementation of acid leaching at a pyrolysis plant, with a biomass capacity of 5 and 50 t h−1, was made and compared with a pyrolysis plant using the untreated biomass. Target products were heating oil and/or additional pyrolytic sugars. It has been calculated that with the leaching step the heat for pyrolysis and drying of the biomass can still be supplied by the combustion of the char and gases, but insufficient excess heat is available to produce electricity for the process. Critical for the economics of the acid leaching pyrolysis process are the amount of extractives in the biomass (organics ending up in the waste water) but not its moisture content. Mechanical dewatering before thermal drying turns out to be very important. The economics of the presented approach turned out to be very sensitive to the plant scale, CAPEX and obviously to the biomass price. At the current market scenario and state of proven techniques the production of sugars and heating oil from bagasse at 50 t h−1 is the most economic option (IRR 15.4%).

Academic research paper on topic "Using pyrolytic acid leaching as a pretreatment step in a biomass fast pyrolysis plant: Process design and economic evaluation"

Biomass and Bioenergy xxx (2016) 1—17

Contents lists available at ScienceDirect

Biomass and Bioenergy

journal homepage: http://www.elsevier.com/locate/biombioe

Research paper

Using pyrolytic acid leaching as a pretreatment step in a biomass fast pyrolysis plant: Process design and economic evaluation

S.R.G. Oudenhoven*, A.G.J. van der Ham, H. van den Berg, R.J.M. Westerhof, S.R.A. Kersten

Sustainable Process Technology Group, Faculty of Science and Technology, University of Twente, Postbus 217, 7500AE Enschede, The Netherlands

ARTICLE INFO

Article history: Received 27 January 2016 Received in revised form 1 July 2016 Accepted 2 July 2016 Available online xxx

Keywords: Fast pyrolysis Acid leaching

Alkali and alkaline earth metals Pyrolytic sugars Process design Techno-economic evaluation

ABSTRACT

Removing alkali and alkaline earth metals (AAEMs) from biomass, with pyrolytic acids, before pyrolysis leads to increased organic oil and sugar yields. These pyrolytic acids are produced and concentrated within the pyrolysis process itself. The purpose of this paper was to evaluate under which conditions acid leaching of pinewood, bagasse and straw can improve the technical and economic feasibility of a py-rolysis process. Therefore, a preliminary process design for the implementation of acid leaching at a pyrolysis plant, with a biomass capacity of 5 and 50 t h-1, was made and compared with a pyrolysis plant using the untreated biomass. Target products were heating oil and/or additional pyrolytic sugars.

It has been calculated that with the leaching step the heat for pyrolysis and drying of the biomass can still be supplied by the combustion of the char and gases, but insufficient excess heat is available to produce electricity for the process. Critical for the economics of the acid leaching pyrolysis process are the amount of extractives in the biomass (organics ending up in the waste water) but not its moisture content. Mechanical dewatering before thermal drying turns out to be very important. The economics of the presented approach turned out to be very sensitive to the plant scale, CAPEX and obviously to the biomass price. At the current market scenario and state of proven techniques the production of sugars and heating oil from bagasse at 50 t h-1 is the most economic option (IRR 15.4%).

© 2016 The Authors. Published by Elsevier Ltd. This is an open access article under the CC BY-NC-ND

license (http://creativecommons.org/licenses/by-nc-nd/4.0/).

1. Introduction

Fast pyrolysis of lignocellulosic biomass is a promising conversion process to depolymerize the biomass building blocks, hemi-cellulose, cellulose and lignin, into a liquid termed bio or pyroly-sis oil. This oil, a mixture of oxygenates and water, can be used directly for heat and power production or further refined to liquid fuels [1—3] and/or chemicals [4]. In this process dry biomass particles are quickly heated to temperatures around 450 °C—550 °C in the absence of oxygen causing the biomass to decompose into gases, vapors (condensed to obtain oil) and char. The oil yield and composition varies largely between different biomass feedstocks [5]. These differences are mainly caused by the varying alkali and alkaline earth metal (AAEM) contents [6], which have been reported to catalyze dehydration reactions leading to increased water and char production [7,8]. The increased water production often leads to phase separation of the oil, which complicates further processing [9]. Moreover potassium and sodium reduce the melting

* Corresponding author. E-mail address: s.r.g.oudenhoven@utwente.nl (S.R.G. Oudenhoven).

temperature of ash leading to problems in the char combustor [10—12]. Typical AAEM concentrations of "clean" debarked wood are around 5 g kg-1 while herbaceous and agricultural waste streams, like straw, contain around 15 g kg-1 [13]. In the authors opinion, the lower organic oil yield and phase separation of the obtained oils from many biomass residue steams high in AAEMs makes these biomass streams not suitable as feedstock for conventional pyrolysis.

Recently we proposed and validated that the majority of AAEMs can be removed from biomass via leaching with organic acids, produced and concentrated within the same pyrolysis process [14]. The different functional blocks required for the proposed process are shown in Fig. 1 [14]. Size reduction of the biomass is required to achieve reasonable AAEM removal rates during acid leaching (discussed in section 3.2). After acid leaching and rinsing the biomass will contain around 75% of moisture by weight, which has to be removed before pyrolysis. The drying is done by a combination of mechanical and thermal treatment. The pyrolytic acids are separated from the majority of the oil (including sugars and phenolics) by applying fractional condensation of the pyrolysis vapors using at least two condensers operated at different temperatures [15]. The

http://dx.doi.org/10.1016/j.biombioe.2016.07.003

0961-9534/© 2016 The Authors. Published by Elsevier Ltd. This is an open access article under the CC BY-NC-ND license (http://creativecommons.org/licenses/by-nc-nd/4.0/).

S.R.G. Oudenhoven et al. / Biomass and Bioenergy xxx (2016) 1—17

Fig. 1. Function block diagram for the pyrolysis of pyrolytic acid leached biomass.

char and gases are combusted to provide the heat required for pyrolysis and drying.

Pyrolytic acid leaching of different biomasses (pinewood, straw, hay and bagasse) reduced the AAEM content to 90 mg kg-1 — 600 mg kg-1 [6]. Pyrolysis of these acid leached biomasses resulted in a large increase of the organic oil yield compared to the untreated biomasses (e.g. straw increased from 370 g kg-1 to 580 g kg-1) [6]. In addition, the selectivity towards anhydrosugars (also referred to as pyrolytic sugars) was largely increased resulting in a high anhydrosugar content in the oils (e.g. bagasse, mass fraction anhydrosugar in 1st condenser increased from 10% to 48% by weight) [6]. These high anhydrosugar concentrations might allow the economic fractionation of the oil into higher economic value products. The sugars can be separated from the oil (including phenolics and furanics) by water extraction followed by ethyl acetate extraction (further discussed in section 3.7). The isolated sugars, obtained from acid leached pinewood, were successfully, after hydrolysis, fermented to bio-ethanol [16] or converted to lipids [17] with comparable yields as obtained from glucose. The residue after sugar separation is rich in phenolics. Tests at laboratory scale showed that this fraction can be used for the production of transportation fuel via hydro deoxygenation (lower H2 consumption and higher C to oil compared to normal pyrolysis oil) [18,19]. To further improve the economic value of the pyrolysis oil, extraction of phenolics is proposed. The extracted phenolics (mainly oligomeric) can be used as feedstock for phenolic resins production [20] or cracked into mono phenolics [21], which would generate the highest value. The product slate including the purification steps, which are often overlooked, can be seen in Fig. 2.

Several studies have evaluated the economics of pyrolysis oil production using untreated biomass. Generally it is found that the biomass price has a large effect on the pyrolysis oil price (~50% for wood) [22,23]. The price of biomass varies a lot for different biomass types e.g. bagasse (~35 $ t-1), empty fruit bunches (15—35 $ t-1), mallee wood (~40 $ t-1) straw (~80 $ t-1) and pinewood (~80 $t-1). It is worthwhile to mention that low cost biomass often has a high ash content. Techno economic studies involving acid leaching as pretreatment step to increase the oil yield and to improve the processability of the oil from high AAEM feedstocks have not been published so far. The economic potential of producing sugars (or ethanol) via pyrolysis has only been studied for two specific cases.

In 1999 So and Brown showed that sugars (for fermentation to ethanol) produced from prehydrolysis, hydrolyzing the hemicel-lulose, followed by fast pyrolysis of the solid residue (cellulose and lignin) had a comparable production cost compared to sugars produced via dilute sulfuric acid hydrolysis or prehydrolysis combined with enzymatic saccharification [24]. In a later paper Brown compared the aforementioned methods for the production of ethanol also with syngas fermentation. It was found that the syngas fermentation route had the lowest ethanol production cost [25]. However, it should be noted that in this study only ethanol was taken as product whereby a large fraction (containing a significant amount of the energy) of the pyrolysis oil remains unused in case of pyrolysis. The second case, the pyrolysis of acid infused biomass (with H2SO4) producing d-glucose (via hydrolysis) and transportation fuels (via hydrogenation using hydrogen produced from the light organics), was evaluated by Zhang et al., in 2013 [26]. An IRR of 11.4% was calculated. It should be noted that the selected method for the production of expensive glucose (600 $ t-1) is rather optimistic, since the sugar stream after hydrolysis will contain next to glucose also sugars produced from the hemicellulose. In addition, the effectiveness of the described acid impregnation method, using sulfuric acid with a concentration of 50% on mass basis, can be questioned because: i) the low amount of liquid (acid) could lead to an uneven acid distribution and therefore not stabilizing all of the AAEMs; and ii) the low pH is expected to lead to dehydration of the biomass.

The objective of this paper is to identify under which conditions acid leaching can improve the technical and economic feasibility of a pyrolysis process. Therefore a process design of a pyrolysis plant with and without acid leaching processing 5 or 50 t of dry biomass per hour was made. The feedstocks studied were: i) pinewood, which has a low AAEM and moisture content and a high feedstock price; ii) straw, which has a high AAEM content, low moisture and a high feedstock price; and iii) bagasse, which has a low AAEM content, high moisture and a low feedstock price. Four different cases, as depicted in Fig. 2, were studied to evaluate the feasibility of producing multiple product streams, these were: i) pyrolysis of untreated biomass to produce heating oil (HOU); ii) pyrolysis of acid leached biomass to produce heating oil (HUAL); iii) pyrolysis of acid leached biomass to produce sugars and heating oil (S&HUAL); and iv) pyrolysis of acid leached biomass to produce sugars, phenolics

Untreated biomass-

Acid leached biomass

Pyrolysis

1 step condensation

Pyrolysis 2 step 1st conderser oil

condensation

Water extraction Ethyl acetate extraction

2" conderser oil To acid leaching

Water insolubles

Hypothetical separation

Product

- Heating oil

• Heating oil ■ Sugars

- Heating oil " Phenolics

• Heating oil

Case HOu

S&P&H0AL

Fig. 2. Overview of the different cases studied.

S.R.G. Oudenhoven et al. / Biomass and Bioenergy xxx (2016) 1—17

and heating oil (S&P&HUAL). The mass and energy balances for these cases were calculated in Aspen Plus©, using our experimental data as input. Based on the obtained results the key factors which influence the economic feasibility for the different cases were identified. Additional benefits and disadvantages (e.g. processabil-ity, process conditions) of the different cases are discussed in the final outlook.

2. Methodology

Based on the functional units shown in Fig. 1 a process design for a pyrolysis plant with and without acid leaching was made. Possible process options for each of the functional units are discussed. As input for the acid leaching process design the removal rate of AAEMs from pinewood and straw was experimentally determined (see section 3.2). The mass and energy balances of untreated and acid leached biomass pyrolysis were calculated using the process flowsheet software Aspen Plus© (V8.6). The input for the pyrolysis process was based on our experimental results obtained from the studied biomass feedstocks [6]. The heat and material balances combined with data from literature for the power consumption of specific equipment, i.e. grinding, pressing and leaching, were used to calculate the operating costs (OPEX). The capital cost (CAPEX) for the pyrolysis plant was based on published data of two commercial plants and a design study. The CAPEX of the acid leaching equipment was obtained from a commercial vendor (De Smet) who has experience with similar technologies. The CAPEX for the sugar separation train was calculated using Aspen Economics©. Since the layout of the phenolics purification was unknown the CAPEX was estimated to be identical to the CAPEX of the sugar separation. Based on the CAPEX and OPEX the economic potential was evaluated based on a net present value (NPV) calculation in Microsoft Excel©. A sensitivity analysis was performed to study the effect of the various parameters, such as feedstock price and product yields, on the economic performance. The methodology in this paper is summarized in Fig. 3.

3. Process design

Fig. 3. Overview of the individual process steps and required data input.

pinewood, straw and bagasse. These biomass feedstocks are delivered as debarked wood chips (moisture mass fraction 15%), straw bales (moisture mass fraction 15%) and wet bagasse fibers (moisture mass fraction 75%). The composition of the feedstocks were determined in our previous study [6] and are shown in Table 1. It can be seen that straw contains the highest AAEM and extractive content.

In this section the possible process options for each of the functional units, as shown in Fig. 1, are discussed. Acid leaching is discussed before grinding since the leaching experiments will provide the optimal biomass particle size and sets thereby the requirements for the grinding. The design of the pyrolysis of acid leached bagasse at a scale of 50 t h-1 is presented in this section (selected as base case), however the pinewood and straw are also discussed. It should be noted that the process design will largely depend on the plant location (integration with other facilities available). In this paper a new standalone plant was assumed, which is likely a worst case scenario. The system boundaries are defined as: i) the plant gate (feedstock at location); ii) products, i.e. pyrolysis oil and/or separated oil fractions (sugars and phenolics); iii) electricity, process water and cooling water are available at the plant location; and iv) waste water can be treated externally conform market price.

3.1. Biomass feedstocks

Fast pyrolysis is generally considered to be capable of processing most dry biomass feedstocks. However the feedstock type and composition will have impacts on the process and economics. For example the feedstock composition will influence the obtained pyrolysis products and the design of the pre-treatment process, e.g. the grinder type, particle size, the leaching apparatus and amount of process water required. The feedstocks studied in this work were

Table 1

Feedstock composition before acid leaching, data from Ref. [6]. Compositional and ultimate analysis are expressed on dry ash free basis (d.a.f.).

Pinewood Straw3 Bagasse

Untreated Untreated Untreated

Compositional analysis, mass fraction in biomass d.a.f. (kg kg 1)

H2O Extractive 0.03 0.06 0

glucose 0.51 0.49 0.44

xylose 0.05 0.25 0.20

galactose 0.02 0.00 0.00

arabinose 0.01 0.03 0.02

mannose 0.12 0.01 0.00

lignin 0.28 0.17 0.23

Ultimate analysis, mass percentage in biomass d.a.f. (% by weight)

C 50.1 48.5 49.3

H 6.1 6.2 6.1

N 0.1 0.4 0.2

Oa 43.7 44.9 44.4

Ash composition, mass concentration in dry biomass (mg kg ')

Na+ 60 117 24

K+ 398 12595 1792

Mg2+ 387 427 402

Ca2+ 1771 2068 700

sum (AAEMs) 2616 15207 2918

Total ash 5260 78151 13387

a Batch 2 in Ref. [6].

S.R.G. Oudenhoven et al. / Biomass and Bioenergy xxx (2016) 1—17

3.2. Acid leaching

The biomass feedstocks studied in this paper are vascular plants containing pores (tracheid's); through which water and nutrients were transported within the plant. The AAEMs exist partly as salts (e.g. oxide, carbonate, oxalate and chloride) deposited in cells or pores after drying [27,28]. The other part of the AAEMs is ionically bound to the lignin or hemicellulose via the empty electron pair of the oxygen atoms in carboxyl groups and hydroxyl groups [28]. Aqueous acid-base reactions, involved in the leaching of the AAEMs, proceed almost instantly [29], therefore it can be expected that the removal of AAEMs is rate limited by mass transfer within the biomass (pores and cell wall). Therefore, the particle size is an important parameter. Small particles will allow fast leaching, resulting in smaller leaching equipment, but demanding a higher electrical energy usage for the grinding. To obtain insights in the removal rate of AAEMs from straw and pinewood with various particle sizes, leaching experiments were performed. To minimize the energy requirement combined with the use of an air closed system it was tested if sufficient AAEM removal rates could be obtained at 30 °C and atmospheric pressure. The experiments and results are presented in supplementary information S1.

Summarizing; the leaching experiments with diluted acetic acid (mass fraction 10%, identical to the 2nd condenser oil) at 30 °C showed that the AAEM removal from ball milled particles (<125 mm, containing no fibrous structure) was almost instantaneous (<2 min) whereas the AAEM removal from cylindrical rods (diameter 27 mm with length of 5,10 or 40 mm) took hours or even days. From these results it can be concluded that the AAEM removal rate is determined by mass transfer in the biomass pores, rather than mass transfer in the cell wall. Leaching with demineralized water at 30 °C removed most of the potassium and sodium (75%— 90% relative to washing with dilute acetic acid) but could remove only 20% of the calcium and magnesium. Although the AAEM content of untreated wheat straw is significantly higher compared to untreated pinewood, the rate of AAEM removal via both water and acid leaching was approximately the same. Based on the experimental results and a 1D diffusion model the particle size for straw and wood was selected to be 2 mm (sieve size) in order to obtain leaching times of 10 min per equilibrium stage. For bagasse it was assumed that no size reduction is required for the acid leaching since the size and structure has been intensively altered before (and

possible during) the sugar extraction to obtain good sugar leaching results.

The AAEM removal equipment should at least consist of an acid leaching step and a rinsing step [14]. The rinsing step is required since biomass retains around 3 times its own dry mass in (leaching or rinsing) liquid, containing dissolved minerals, after draining. To reduce the required amount of process water the rinsing should be performed in countercurrent operation. Fig. 4 shows a schematic overview of the leaching and rinsing process. In this process the 2nd condenser oil, containing acids (see Fig. 2), is added in the first stage to directly dissolve all AAEMs, the later stages are used for the rinsing. For some high AAEM biomass feedstocks, the amount of acetic acid produced in the pyrolysis process may not be sufficient to dissolve both the mono and divalent AAEMs [6]. Therefore, a pre-water wash in order to dissolve the water soluble monovalent AAEMs before acid leaching is required.

In practice there are two options for contacting the biomass with the (leaching or rinsing) liquid, these are: i) moving the liquid through the "fixed" particles, which is termed percolation; and ii) the biomass particles are added to the liquid for a desired time and subsequently recovered. Possible process equipment to carrier out the leaching of biomass include; Bollman-type extractors, rotocel extractors belt percolator (e.g. sugar cane diffuser from De Smet), Kennedy extractors and Bonotto extractor [30]. A diffuser, as applied industrially to extract sugar from sugar cane, was selected because of its simple design, low energy consumption, adaptable to variations in feed capacity and little amount of damage to the fibers (leading to fines). If larger particles are used the AAEM removal rate can be increased by using ultrasonic mixing [31], mechanical pressing of the biomass (forced liquid flow in/out particle) [32] or increasing the leaching temperature (requiring a closed system due to increased volatility of the acids).

The process shown in Fig. 4 can be mathematically described as countercurrent liquid-liquid extraction. The leaching liquid remaining in the biomass pores, containing the dissolved AAEMs, is regarded as the feed to the next stage. The free rinsing water, outside the biomass, is regarded as the solvent. Since the solubility of the AAEM acetate salts is much higher than the concentrations obtained during leaching the formation of salts can be neglected. If equilibrium between the water inside and outside of the biomass is assumed (sufficient contact time) this system can be described by the following equations (condenser 2 liquid is neglected since flow is much smaller than rinsing water)

Stage 1 : 42/1*C2 - (41/2 + 41/waste)*C1 = - 4 biomass,dry* Cbiomass,dry

For 2 till n - 1: 4i-1/i*Ci-1 + 4i+1/i*Ci+1 - (4i/i-1 + 4i/i+1)*C = 0

Stage n : 4n-1/n*Cn-1 - i 4n/n-1 + 4n/dryer) *Cn = -4fresh water*Cfresh water

S.R.G. Oudenhoven et al. | Biomass and Bioenergy xxx (20Í6) Í—Í?

Rinsing 2nd Condenser oil Water (containingacids)

Rinsing Water

Grinded biomass

Optional prewash _ t o_r_e m_o_v e K_ a n d_ Na_ _

I— Í- Í I

■ 2 n-1 > n

l_l _1 _1 _1

Waste water Waste water

Fig. 4. Schematic overview of a countercurrent extractor for leaching and rinsing.

In these equations 4 is the mass flow of both water streams (remaining within the biomass pores or free rinsing water) or dry biomass and C is the concentration of AAEMs in water or dry biomass. The mass flow of water (remaining in the biomass pores) to the next stage (i / i + 1) is equal to 3 times the mass flow of biomass. Using this set of equation together with a mass balance of the water for each stage, the required amount of rinsing water as function of the number of equilibrium stages required to reduce the AAEM content to a specific concentration given an AAEM concentration of the process (rinsing) water can be calculated. The equations were written as a matrix and solved in Matlab©. The composition of process water largely depend on the location [33]. In this study we assume that the process water available at the location has an AAEM content of 50 mg kg-1. Fig. 5a shows the amount of equilibrium stages required to reduce the AAEM content of pinewood and bagasse to 200 mg kg-1 using process water containing an AAEM concentration of 50 mg kg-1 water. Fig. 5b shows the results for straw which requires a pre-water wash, in which the potassium and sodium content is reduced by 80% (90% for straw was possible see Fig. S5), and the acid leaching step, in which the AAEM concentration is reduced to 200 mg kg-1. These figures shows that the amount of water required for acid leaching flattened out after 4 to 5 stages. For the straw water wash the amount of water required becomes quite stable after 3 to 4 stages. Therefore the amount of equilibrium stages required for acid leaching will be set to 5 and for the pre-water wash to 3. Moreover it can be seen that around 4, 7.5 (sum of water wash and acid leaching) and 3 kg of rinsing water will be requires per kg of dry biomass for pinewood, straw and bagasse, respectively. Note, the amount of "fresh" process water required can be reduced by recycling - the water from the biomass drying.

3.3. Grinding

As mentioned in the previous section (3.2) the biomass particle size was selected to be 2 mm to achieve residence times of 10 min per equilibrium stage (resulting in a total residence time of 50 min in the diffuser). It is worthwhile to mention that several pyrolysis reactor types like (bubbling/circulating) fluidized bed require small particles. The selected particle sizes are also suitable for these types of pyrolysis reactors so no further size reduction is required. Therefore, this particle size is also selected for the untreated pinewood and straw case. The received straw bales first need to be broken up where after the straw is milled down to the required particle size. Literature is not clear what device (hammer or knife mill) is most favorable in terms of low amount of fines and low energy consumption [34,35]. For the wood chips a hammer mill (screen size 6 mm) will be used (based on [36,37]). It is worth mentioning that the grinding energy for pinewood (50 kWh t-1) is significantly higher compared to straw (15 kWh t-1). Depending on the impurities, e.g. stones or metal nails, in the received feedstocks a (metal) separator can be installed before the grinder to prevent damage of the equipment. For the bagasse no size reduction is required, however it is assumed that the cane has been depithed, thereby removing the fines and majority of AAEMs.

3.4. Dewatering and drying

The biomass leaving the leaching process will contain a moisture content of ~75% on mass basis. Removing this amount of water via evaporation will consume a large fraction of the energy present in the biomass (only heat of evaporation per kg wet bagasse is 1.8 MJ kgwhile the HHV of the wet bagasse is ~4.8 MJ kg-1). Therefore mechanical dewatering, as applied industrial after the sugar cane diffuser, will be applied as a first step to remove the majority of water. The mechanical drying of bagasse is widely applied. Typical moisture contents ~46% on mass fraction after pressing are reported [38,39]. Unfortunately, for straw and wood fibers (saw dust) less information is available. Nevertheless, comparable results are expected. Typical equipment used to mechanically dewater biomass include centrifuges, belt presses, screw presses and roller press [30].

Westerhof et al. showed that a moisture content up to 20% on mass fraction has no significant effect on the organic oil yield during fast pyrolysis in a fluidized bed [40]. Reducing (or

E 14 H □

£•12 T3

cr 6 £

- Pine wood

--•■--• --•--

2 4 6 8 10

Number of equilibrium stages (-)

^15 H £ er

2 10 i— 0

Ьг „.

■ Straw water wash (removal of 80 % monovalent) • Straw acid leaching (reduction to 200 mg kg ~1)

2 4 6 8

Number of equilibrium stages (-)

Fig. 5. Amount of rinsing water required versus the number of stages to reduce the AAEM content in dried biomass 200 mg kg 1 (assuming sufficient acid is present in the 1st stage). In the model the pinewood and straw (to water wash) contained 15% moisture by weight. The bagasse and straw (to acid leaching) contained 75% moisture by weight.

S.R.G. Oudenhoven et al. / Biomass and Bioenergy xxx (2016) 1—17

2nd condenser oil (~10 % acetic acid 1.651 h1) From Fig. 7

Fresh water (301 h1)

Bagasse

(after sugar extraction)

Steam (140°C 3 bara 471 h"1)

Waste water (1831 h1)

Bagasse (140°C 501 h"1) r—i To Fig. 7

Steam L--1

(530°C 3 bara Steam 1521 h1) (140°C 3.5 bara From Fig. 7 1501 h1) To Fig.7

Fig. 6. Schematic overview of leaching and drying bagasse at 50 t h 1 Stream sizes and temperatures are based on mass and energy balance (discussed in section 4).

completely removing) the moisture before pyrolysis will: i) reduce the required temperature of the used energy carrier since the reactor is operated at a higher temperature then the dryer; ii) reduce the volumetric vapor flow from the reactor, thereby decreasing the size of cyclones and condensers; and iii) increases the organics content of the 2nd condenser oil, which is beneficial for the acid concentration and its heating value (see section 3.8). Therefore it was decided to dry the biomass completely before pyrolysis.

The thermal energy required for drying can be supplied via hot gas (e.g. flue gas) or steam [41]. Because of: i) the low self-ignition temperature of dry biomass (fines), ii) the high moisture content of the feed (requiring large amounts of heat), and iii) possibility to utilize the condensation heat at higher temperatures a superheated steam dryer has been selected [42]. To reduce the energy requirement of the compressor in the steam cycle and the size of equipment (i.e. steam pipes in combustor and the dryer) the dryer will be operated at 3 bar and 140 °C. Note that higher pressures will increase the drying temperature. See Fig. 6 for a schematic representation of the pretreatment section for 50 t h-1 bagasse. For the "dry" feedstocks (untreated pinewood and straw, containing 15% moisture) significant less energy is required for drying. As a result, it is less important to recover the evaporation energy. Moreover the mass and energy balance showed that the flue gas from the

combustor (after heating the recycle gas fed to the pyrolysis reactor) still contains enough energy to dry the biomass (this is not the case for the wet bagasse and acid leached feedstocks). Therefore for untreated pinewood and untreated straw it was chosen to use flue gas drying instead of steam.

3.5. Fast pyrolysis and char combustor

Several pyrolysis reactor configurations have been studied and implemented in industry. These configurations include; bubbling fluidized bed (Dynamotive), circulating fluidized bed (Ensyn and Fortrum), rotating cone reactor (BTG), cyclone reactor (Latvian State Institute of Wood chemistry) and auger reactor (Lurgi together with Karlsruhe and Abritech) [43,44]. The energy required for pyrolysis (heating the biomass, reaction enthalpy and heat loss of the reactor) per kg of dry biomass is estimated to be in the range of 1.2—2.5 MJ kg-1 of dry matter [22,45,46]. Based on typical yields and composition of the char and gases from the different untreated biomass feedstocks it can be calculated that burning the char provides on biomass d.a.f. (0.15 kg kg-1 * 32 MJ kg-1) ~ 4.8 MJ kg-1 and burning the produced gases produces on biomass d.a.f. (0.3 kg kg-1 * 8 MJ kg-1) ~2.4 MJ kg-1. It should be noticed that after acid leaching the char yield largely decreases and becomes between 0.054 kg kg-1 and 0.086 kg kg-1 (see Table 2) [6], which

Table 2

Pyrolysis products from untreated and acid leached biomass in kg kg-1. Part of the data from Ref. [6].

Compound Pinewood Straw Bagasse

Untreated Acid leached Untreated Acid leached Untreated Acid leached

Produced water 0.117 0.085 0.113 0.077 0.104 0.044

Char 0.100 0.086 0.136 0.072 0.080 0.054

CO 0.140 0.130 0.149 0.132 0.162 0.120

CO2 0.116 0.088 0.200 0.070 0.247 0.125

CH4 0.020 0.020 0.020 0.015 0.038 0.020

C2+ 0.010 0.010 0.010 0.010 0.010 0.010

Levoglucosan 0.044 0.155 0.007 0.157 0.020 0.200

Mannosan 0.012 0.030 0 0.001 0.002 0

Xylosan 0 0.010 0.002 0.030 0 0.043

Furanics 0.020 0.020 0.020 0.020 0.020 0.020

monomeric phenols 0.020 0.020 0.020 0.020 0.020 0.020

Water soluble phenolics 0.038 0.058 0.043 0.078 0.047 0.100

Water insoluble 0.122 0.102 0.054 0.043 0.093 0.041

Unknown cond 1 0.102 0.074 0.064 0.106 0.032 0.095

acetic acid 0.025 0.018 0.018 0.019 0.025 0.018

formic acid 0.010 0.010 0.010 0.010 0.010 0.010

Lights 0.104 0.084 0.134 0.140 0.090 0.080

S.R.G. Oudenhoven et al. / Biomass and Bioenergy xxx (2016) 1—17 7

corresponds to 1.8 MJ kg-1—2.8 MJ kg-1. Moreover in case of acid leaching a significant amount of heat is required to dry the wet biomass after leaching. Based on these numbers it can be concluded that for the pyrolysis of untreated biomass the produced gases and char contain significant more energy than the heat required for the pyrolysis process. Therefore the excess heat is often converted into electricity, which is partially used to provide the electricity consumption of the pyrolysis plant. The energy balance for acid leached biomass is presented and discussed in 4.1.

The heat to the reactor can be supplied via: i) a solid heat carrier (e.g. sand or steel beets), ii) hot gas, iii) superheated steam, iv) via the reactor wall, v) intermittent reactor concept (operation switching between pyrolysis and combustion mode) [47], or vi) via (fired) heating tubes. Pyrolysis reactions occur very fast (seconds) at high temperature (>500 °C) [48] and are therefore often limited by the heat supply. Supplying the heat via heat (fired) tubes or the reactor wall will required a large heat exchange area. Therefore reactors heated via the wall will be limited in size. Pernikes et al. reported superheated steam to biomass ratios between 2 kg kg-1and 7 kg kg-1 [49]. It can be calculated that the energy required for producing this amount of superheated steam is larger than the amount of energy in the produced char and gases. Supplying the process heat via recycling (fluidization) gases will results in large gas streams (heating recycle gas from 30 °C to 530 °C will cost ~0.5 MJ kg-1) and requires quite high gas temperatures. Moreover feeding steam or large gas streams to the reactor will results in larger cyclones and condensation equipment. The intermittent pyrolysis reactor concept as proposed by Siemons and Baaijens is an interesting method [47], however this concept is still in its early stages. Therefore the use of a heat carrier was selected in this study.

The goal of this study was to evaluate under which conditions acid leaching can improve the technical and economic feasibility of the pyrolysis process. So far none of the pyrolysis reactor concepts is outperforming the others, so in this study we don't want to select

a specific design (other than heating via a solid heat carrier). In this light, the input for the Aspen Plus© simulation was based on the results, summarized in Table 2, obtained from the studied biomass feedstocks in our fluidized bed setup. Therefore, the process conditions (pyrolysis temperature 530 °C) were chosen identical to the experimental conditions used in our previous paper [6]. We want to emphasize that these are not necessarily the optimal pyrolysis conditions. In the reactor the biomass is converted into a mixture of vapors, gases and char. A series of cyclones separates the particles from the gases and vapors. Part of the pyrolysis gases are recycled back to the pyrolysis reactor for fluidization or removing the py-rolysis vapors from the hot reactor (recycle gas flow on biomass d.a.f. 1 kg kg-1, based on our setup [6]). The char and surplus of pyrolysis gases are burned in the combustor to provide the heat for the circulating heat carrier. The combustor is operated at 580 °C, thus 50 °C above the reactor temperature. Steam is produced from the excess heat of the combustor and used for biomass drying and electricity production (in case of untreated pinewood and straw). Air is fed at an air to fuel ratio (l) of 1.3]. See Fig. 7 for a schematic representation of the combustor and pyrolysis reactor.

3.6. Oil collection

Fractional condensation of the pyrolysis vapors is used as an inexpensive first downstream separation technique to separate the acids, required for acid leaching, from the sugars and phenolics (products). Pyrolysis oil collection systems described in the literature usually consist of a combination of several: i) spray condensers [40], ii) electrostatic precipitators (ESP) [50], or iii) (shell and tube) heat exchangers [51,52]. Inside a condenser the vapor stream is quenched, causing the vapors to condense and thus forming aerosols. Spray condensers [40] and ESPs [50] have proven to be very efficient in collecting these aerosols. Therefore, the vapors and gases leaving the condenser are at vapor-liquid equilibrium, allowing controlled fractionation, as demonstrated by Westerhof

2nd condenser liquid to furnace (4.2 thr1)

Fig. 7. Schematic overview of the pyrolysis, combustion and oil recovery section for the pyrolysis of acid leached bagasse. Stream sizes and temperatures are based on the mass and energy balance from Aspen Plus© (discussed in section 4).

S.R.G. Oudenhoven et al. / Biomass and Bioenergy xxx (2016) 1—17

et al. [40]. Based on the simplicity and better heat transfer of the spray condensers this type has been selected. A two stage condensation system was selected in this study. Although, more condenser in series (high temperature condenser to collect the water insolubles) might be an option to improve the fractionation of the oil even further (as discussed in 3.7). The temperature of the condensers is controlled by the recirculated bio-oil. In the first condenser the temperature of the recirculated oil is controlled via indirect cooling (possibility to recover heat at temperatures below condenser temperature) while the oil in the second condenser is cooled indirect with cooling water. The recovery of heat from the pyrolysis vapors, before entering the condensers, is not included due to the expected fouling caused by condensation inside the heat exchangers.

3.7. Sugar and aromatic separation

The 1st condenser oil, rich in sugars and phenolic compounds, needs to be separated when applied as feedstock for fermentation (sugars), monomeric phenolic or phenolic resins (phenolic fraction) production. Pyrolysis oil separates into an aqueous phase, containing oxygenates including the anhydrosugars, and a tarry organic phase, containing mainly phenolic oligomers [53,54], when the water content reaches 30%—35% on mass basis [55]. This principle is often applied as a first step in the separation of pyrolysis oil to remove the "sticky" water insolubles [16,18,19,21,56,57]. Rover et al. studied the phase separation of fractional condensed pyrolysis oils produced from acid infused red oak to recover a sugar-rich solution (for fermentation) and a phenolic oligomers stream [57]. It was shown that the sugars can be extracted effectively (>93% mass basis) by two successive water extractions using an oil to water ratio of 1:1 (mass basis). However, fermentation inhibitors, such as 5—hydroxymethylfurfural, guaiacol and phenol, were also extracted leading to suppressed growth of the used E. coli (biocat-alyst for ethanol). Removal of the inhibitors was performed by overliming (with NaOH, Ca(OH)2 and NH4OH), however the concentration of the sugars in the fermentation broth was still limited to maximal 2% mass fraction [57]. Moreover overliming leads to the formation of a solid waste stream which needs to be deposited. Other options to remove these inhibitors (mainly aromatics) is extraction with ethyl acetate. Lugue et al. showed that the aqueous phase after ethyl acetate extraction and hydrolysis could be successfully fermented to ethanol [16]. However this method has several drawbacks including: i) the solubility of ethyl acetate in water is around 81 g dm~3 [58] requiring the recovery of ethyl acetate from the aqueous stream; ii) ethyl acetate and water form an azeotrope, however the recycled ethyl acetate is reused for the extraction of the aqueous phase from the water separation therefore this azeotrope does not cause any practical problems; and iii) ethyl acetate is more volatile than the extracted aromatics which requires the evaporation of the solvent [58]. Another options for oil separation is by crystallization of the sugars in cold acetone as applied at the pyrolysis plant located at the Krasnodar hydrolysis plant (USSR) operated between 1958 and 1970 [59]. Unfortunately this method is quite sensitive to the pyrolysis oil composition; high water contents (<10% mass fraction [49]) and the presence of water insoluble compounds [49] results in poor separation results. Therefore, this method was only applied for the pyrolysis of hy-drolyzed lignocellulose at mild temperature (<420 °C using steam as heat carrier) combined with fractional condensation of the vapors (T 1st condenser < 100 °C [59]).

Because of the good fermentation results (at a high concentration) after ethyl acetate extraction combined with the presence of water insolubles in the produced oils the separation of sugars from the 1st condenser oil will be calculated based on the ethyl acetate

extraction process, even though it is far from ideal. Fig. 8 shows a schematic representation of the separation train for the 1st condenser oil from acid leached bagasse. In the process the 1st condenser oil is first extracted with water (water to feed ratio 1:1) in a countercurrent extractor with 3 stages. The aqueous stream is hereafter extracted with ethyl acetate (amount of ethyl acetate per aqueous phase 0.2 kg kg-1) in a countercurrent extractor with 4 stages. The ethyl acetate extract is separated in a distillation column (6 stages). The ethyl acetate dissolved in the aqueous stream is recovered in a second distillation column (8 stages). Note that the oil composition between the different acid leaching cases varies (e.g. water insoluble yield for bagasse 2 t h-1 versus pinewood 5 t h-1). The stream sizes were estimated based on: i) our experimental experience for the phase separation of the 1st condenser oil; ii) the liquid-liquid equilibrium data determined by Li et al. [58] for the model mixture of ethyl-acetate, water, levoglucosan and guaiacol [58]; and, Aspen Plus© simulation using the data Li et al. [58] for the CAPEX and OPEX of the distillation columns.

The aromatic fraction from the aforementioned separation (water insolubles together with ethyl acetate extract) still contains non phenolic compounds (furanics and lumped unknown compounds). So far no research has been performed to further purify/ extract the phenolic compounds from this fraction. Therefore in the S&P&HOAL cases this separation will be treated as a hypothetical separation. The CAPEX and OPEX are assumed to be identical to that of the sugar extraction from the 1st condenser oil. The remaining organics will be used as heating oil.

3.8. Waste water

The waste water from the leaching process (see Fig. 4) will contain: i) dissolved light oxygenates (from the 2nd condenser liquid), ii) organic extractives (from the biomass), and iii) minerals (from the biomass). This stream needs to be treated before being discharged or reused. Several options to treat the dissolved or-ganics have been studied, these include: i) aerobic bacterial treatment (rather costly), ii) reverse osmose [60,61], iii) anaerobic digestion to methane [62,63], iv) fermentation to lipids [64], v), and v) microbial electrolysis to hydrogen [65]. Unfortunately, the aforementioned techniques are still in initial research scale except for aerobic bacterial treatment. Therefor the cost (330 $ t-1 of organic material) for removing the organics via aerobic treatment (worst case) was used in the economic evaluation, however the potential of the other techniques will be discussed. To reduce the amount of organics in the waste water stream from the biomass leaching a fraction (depending on the required amount of acid required for leaching) of the 2nd condenser oil is fed to the combustor (see Fig. 7). The dissolved minerals, mainly potassium and calcium, in the process water can be recovered via reverse osmosis combined with evaporation of the brine. However, since the sodium adsorption ratio of the leaching liquid is rater low the most favorable application would be to us it as irrigation water, to compensate for the nutrients taken from the soil to prevent soil depletion. The use as irrigation water, after organic treatment, is selected for the base case. Note that the sludge from the aerobic treatment will contain a fraction of the dissolved minerals.

4. Process evaluation: mass and energy balances

Based on the technologies discussed in the previous section, flowsheets for pyrolyzing untreated and acid leached bagasse were made. The flowsheet for the pyrolysis of acid leached bagasse is visualized in Figs. 6 and 7. The flowsheets for acid leached pine-wood and straw have an almost identical layout; a grinder is required before leaching and the exact mass flows (including 2nd

S.R.G. Oudenhoven et al. / Biomass and Bioenergy xxx (2016) 1—17

Fig. 8. Schematic overview of the oil separation train used to purify the sugars in the 1st condenser oil from the pyrolysis of acid leached bagasse at 50 t h

Table 3

Evaluation of mass and energy balances for the pyrolysis of untreated and acid leached feedstocks (HOU and HOAL cases). Yields on biomass dry ash free (d.a.f). Note in case of untreated biomass the whole oil is the product whereas for acid leached biomass the product is only the 1st condenser oil.

Pinewood Straw Bagasse

Untreated Acid leached Untreated Acid leached Untreated Acid leached

Mass to oila (kg kg 1) 0.61 0.48 0.44 0.44 0.46 0.54

C to oila (kg kg-1) 0.54 0.55 0.32 0.48 0.37 0.56

Energy to oila (kg kg-1) 0.54 0.52 0.33 0.51 0.40 0.57

Mass to sugar (kg kg-1) 0.06 0.19 0.01 0.16 0.02 0.24

C to sugar (kg kg-1) 0.05 0.17 0.01 0.15 0.02 0.22

Electricity added per energy in feed (MJ MJ-1) 0.02 0.03 0.01 0.02 0.02 0.02

a Oil in case of untreated whole oil in case of acid leached only 1st condenser fraction.

condenser liquid for leaching) and temperatures of the streams differ slightly. The flowsheets for the pyrolysis of untreated pine-wood and straw are shown in Figs S8 and S9, and for untreated bagasse in Figs S10 and S11. These flowsheets were translated into several Aspen Plus© models to calculated the mass, carbon and energy balances. The model input for the pyrolysis process was based on the pyrolysis yields, summarized in Table 2, obtained from the studied biomass feedstocks pyrolyzed in our fluidized bed reactor at 530 °C [6]. Other input to the model and assumptions are explained in Section 10.3 (supplement). The solubility of AAEMs, as

function of the amount of acetic acid added, and the product distribution over both condensers, as function of the condenser temperature, as calculated with Aspen Plus© were compared with experimental results to validate the used models (results presented in supplement S10.4.4). The agreement was satisfactory.

4.1. Overall mass, carbon and energy balances

Table 3 summarizes the mass and energy balances for the HOu and HOAL cases (see Fig. 2) for the different feedstocks (S&HOAL and

S.R.G. Oudenhoven et al. / Biomass and Bioenergy xxx (2016) 1—17

S&P&HOAL only contain additional product separation compared to HOAL). After acid leaching the mass conversion to oil only increased for bagasse. However the carbon and energy efficiency to oil also increased for straw after acid leaching. The decreased yield in oil for pinewood can be explained by the rater small amount of additional oil after acid leaching (due to the low AAEM content in untreated) combined with the loss of the 2nd condenser oil. For all biomasses a large increase in the mass and carbon efficiency from biomass to sugar was obtained after acid leaching. As expected the yields from the different feedstocks were quite comparable after acid leaching. The electricity usage per amount of thermal energy in the biomass varied between 0.01 and 0.03 MJ MJ-1. The differences in the electricity usage can be attributed to: i) the additional power consumption of the diffuser and press for acid leaching, and ii) the grinding of pinewood requires more power than straw. Based on the Aspen Plus© simulation it was quantified that for all biomass feeds the heat for both pyrolysis and biomass drying can be supplied by combustion of the char and gases. In case of the pyrolysis of untreated pinewood and straw a large amount of excess heat is available which can be converted to produce electricity (more than required by the process). All relevant mass and energy flows for all cases are summarized in Tables S6 to S11 in supplementary information.

4.2. Detailed mass, carbon and energy balance for the pyrolysis of acid leached bagasse

Fig. 9 shows the mass flows for the pyrolysis of acid leached bagasse at a feed capacity of 50 t h-1 dry mass. Note, the superheated steam stream from the combustor to the dryer and vice versa (red dotted arrow) is not displayed at the correct scale for better readability. It can be seen that the majority of the initial dry

biomass leaves the process as 1st condenser oil (54%) or flue gas (41%). Only a small amount of the organic biomass leaves the process via the waste water stream. Because of the superheated steam drying, the moisture remaining after the dewatering press (47 t h-1) is converted into low pressure steam (140 ° C at 3 bar). This steam is in the current design unused, but might be (partly) used for: i) the distillation columns to separate the sugars from the ethyl acetate (Fig 8), ii) to evaporate the brine obtained after reverse osmosis of the waste water in case the inorganic containing water cannot be discharged, or iii) to crystalize sugar when the process is located at a sugar cane mill.

Fig. 10 shows the carbon mass balance (the recycle gas stream to the reactor was excluded for better readability). Bagasse consist for approximately 50% on mass basis of carbon (Table 1). It can be seen that 56% (untreated bagasse 37%) of the carbon entering the process is recovered as product in the 1st condenser. The carbon balance shows that the majority of the unrecovered carbon is lost to the gas (20%) and char stream 10%). The char and gases are burned in the combustor to provide the heat for the pyrolysis reactor and biomass drying.

Fig. 11 shows the energy flows for the pyrolysis of acid leached bagasse. The energy content of the organic streams was calculated using the Milne equation to obtain the higher heating values (HHV) of the streams. From this figure it can be seen that a part of the energy is lost during the condensation of the vapors (1st condenser 8% and 2nd condenser 1%), which can be recovered at the respective temperature of the condenser (90 °C and 25 °C). Moreover, the steam produced contains 12% of the energy in the feed. In the end 57% (untreated bagasse 40%) of the energy in the feed is recovered in the 1st condenser oil (product). Based on Fig. 11 it can be concluded that the overall energy recovery (57%) in the final product for the pyrolysis of acid leached biomass cannot be

Combustor (Heat, steam and power Generation)

Fig. 9. Mass balance for the pyrolysis of acid leached bagasse at 50 t h 1 scale. Steam stream from combustor to dryer and back is not represented at correct scale for better readability.

S.R.G. Oudenhoven et al. / Biomass and Bioenergy xxx (2016) 1—17 11

Fig. 10. Carbon mass balance for the pyrolysis of acid leached bagasse at 50 t h 1 scale. The gas recycle to the reactor is excluded for better readability (this doesn't change the outcome). Note elemental composition based on experimental determined product composition (see Table S5).

Steam 140°C at 3 bar

2nd Cond oil to leaching 26 GJ h"1

Char 94 GJ h'1

Compressor Heat reactor steam

1710 kW 0

Coolant 1st condenser ~85 "C

_ nser

Coolant 2nd condenser ~25-C

115 SJh 2»d tond oil to combustor 68GJI11

Ia cond oil 540 GJ h'1

Air blower 1500 kW

Fig. 11. Energy flows for the pyrolysis of acid leached bagasse at 50 t h-1 scale. Energy content of the organic streams was based on HHV. Steam stream from combustor to dryer represents only the heat added to drying process.

increased much further without using an external heat sources (from wood, coal natural gas or fraction of produced oil). The simulation proved, as expected, that for the pyrolysis of untreated biomasses a larger fraction of the energy is recovered in the char and gases, which can be used for steam or electricity production. The total power requirement for the plant, including grinding, acid leaching pressing etc., is calculated to be 5.8 MW (bagasse 50 t h-1

HOAL), which in case of dry (15% moisture) untreated biomass can be produced from the excess heat produced in the combustor.

4.3. Flows and composition of the waste water streams

The sizes and composition of the waste water streams for the different biomasses are summarized in Table 4. The amount of

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Table 4

Amount and composition of the waste water streams for pyrolysis of acid leached biomass at a feedstock capacity of 50 t h_1.

Pinewood Straw Bagasse

Waste water from prewash (t h-1) — 107 —

Organic content (% mass fraction) — 2.4 —

Estimated COD (g kg-1) — 40.5 —

Inorganic content (% mass fraction) — 0.8 —

Waste water from acid leaching (t h-1) 91 187 183

Organic content (% mass fraction) 2.5 2.7 0.4

Estimated COD (g kg-1) 58 70 13

Inorganic content (% mass fraction) 0.5 0.4 0.2

waste water per kg of dry biomass was 1.8, 5.9 and 3.7 kg kg-1 for pinewood, straw and bagasse respectively. The large water stream for straw was caused by the mandatory separation of the two waste stream (see Fig. 4), to prevent precipitation of the calcium during the water prewash. The difference in the waste water stream flow between pinewood and bagasse was caused by the large amount of water in the bagasse feed, which is fed at the first stage. The differences in the stream composition were mainly caused by both the higher extractive and AAEM content in straw. As a result of the higher AAEM content more 2nd condenser oil is required for the acid leaching. The estimated Chemical oxygen demand (COD) of the waste water streams varied between 13 and 70 g kg-1. The biological oxygen demand (BOD), which can only be experimentally determined but often estimated as 70% of COD [66], is a good indicator for the potential to produce methane from water rich waste streams via anaerobic digestion. At BOD concentrations over 1 g kg-1 anaerobic digestion before aerobic bacterial treatment is usually economic attractive [66]. It can be seen that all waste water streams have a higher COD. Therefore, it is expected that implementing anaerobic digestion before the aerobic bacterial treatment can reduce the costs associated with the waste water treatment (not done in this study). A good indication of the required waste water treatment layout and associate costs can be obtained from the design studies by Merrick & Company [66] and Brown & Caldwell [67], prepared for the biomass to ethanol process studied by NREL [68]. It should be noticed that the water streams from the ethanol process contains high ammonia concentrations which is not the case for waste water streams from the pyrolysis process.

5. Economic assessment

In this section the economic feasibility of the processes, as designed in the previous sections, is evaluated. The economic potential was evaluated using a net present value (NPV) calculation based on the CAPEX and OPEX.

5.1. CAPEX and OPEX

The capital cost (CAPEX) of the pyrolysis plant (including storage, grinding, pyrolysis and combustion) was based on published data for various pyrolysis systems. This data is summarized in Table S12. Interestingly, the capital cost for the different reactor

Table 5

CAPEX for the different cases at 50 t h-1 feedstock capacity (2014 basis).

Cases Total CAPEX (million $)

HOu 121

HOal 165

S&HOal 171

S&P&HOAL 178

Table 6

Material and utilities costs used in the economic evaluation.

Materials and utilities Price Source

Pinewood chips 80 $ t-1 of dry matter [69]

Straw bales 80 $ t-1 of dry matter

Bagasse 35 $ t-1 (bagasse) of dry matter [70]

Electricity 0.06 $ per KWh [71]

Process water 0.2 $ m3 [71]

Water treatment 330 $ t-1 of organic [71]

Solid disposal (landfill) 170 $ t-1 of dry material [71]

types, i.e. rotating cone, circulating fluidized bed and auger, did not differ much. An estimate of the CAPEX of the diffuser (for acid leaching) was supplied by De Smet, this data can be found in Table S13. The CAPEX for the separation train required to extract the sugars from the 1st condenser oil was estimated using Aspen Economics© (see Table S14). For the hypothetical purification of the aromatic fraction an identical CAPEX as the sugar separation train was assumed. Table 5 summarizes the total CAPEX for the different cases. Based on these numbers it can be calculated that implementing acid leaching at a new pyrolysis plant will increase the total CAPEX by 36% (HOAL), 41% (S&HOAL) or 47% (S&P&HOAL).

Tables 6 and 7 show the costs of materials and utilities and the prices for the different products respectively. Producing sugars for fermentation (350 $ t-1) will increase the economic value of this oil fraction since sugar has a HHV of 17.5 MJ kg-1, which corresponds to 210 $ t-1 (at 12 $ Gj-1). The HHV of the phenolics (based on the elemental composition of the water insolubles) is around 28 MJ kg-1, which corresponds to 330 $ t-1. The price for the phenolic fraction was estimated (400 $ t-1) since no large batches of this fraction have been made and therefore no industrial applications currently exist.

5.2. Economic feasibility of different cases and feedstocks

The annual operating costs and earnings for the different cases at 501 h-1 are shown in Fig. 12. Form this figure it can be seen that:

• None of the cases using pinewood or straw as feedstock results in a profit while for bagasse in all cases a profit is obtained. This difference is mainly caused by the higher biomass costs for pinewood and straw.

• The waste water treatment, converting the dissolved organics, delivers a huge cost in case of straw, which is caused by both the higher extractive content and AAEM content (requiring more 2nd condenser liquid for leaching) in straw.

• No increase in annual income is obtained after acid leaching of pinewood when only heating oil is targeted. This was expected based on the energy to oil efficiency for pinewood (Table 3)

• The different feedstocks show mutual trends after acid leaching, the differences are mainly caused by: i) the larger amount of extractives in the straw, and ii) higher lignin content in pine-wood yielding more phenolics.

• The production of sugars, phenolics and heating oil (S&P&HOAL) from acid leached biomass generates the largest annual income, however compared to producing only sugars and heating oil (S&HOAL) the increase in earnings was rather small. Recovering only sugars and phenolics (no heating oil) from the oil (see S&P&HOAL case) generates a lower annual earning then recovering sugars and heating oil (S&HOAL).

Fig. 13 shows the annual cash flows for the different cases at a feedstock capacity of 5 t h-1. It can be seen that at this scale the capital, fixed (fraction of CAPEX) and labor cost are too large to

S.R.G. Oudenhoven et al. I Biomass and Bioenergy xxx (20Í6) Í—Í7

Table 7

Product prices used in the economic evaluation.

Products

Source

Heating oil

Sugars for fermentation

Aromatics

12 S Gj-1 (on HHV basis) 350 S t-1 of dry sugar 400 S t-1 of dry aromatics 6.6 S t-1

Average (5 year) fuel oil #5 Based on sugar #11 Own estimate [71]

ID 80 О

Л. 70

Ш |60

!» a>

Costs: B Biomass B Fresh water U Electricity (consumption) B Waste water treatment Ash disposal U Capital depriciatlon B Labour B Fixed costs

Earnings:

Heating oil Sugars Aromatics U Electricity (production) H steam

«а со

Pinewood

О X öS CL об СО

«а со

°8 О.

ов И

Fig. 12. Annual operating costs and earnings for the different cases at 501 h 1 scale. For each individual case the first column shows the operating costs and the second column the earnings.

Costs:

U Biomass

H Fresh water U Electricity (consumption) U Waste water treatment Ash disposal B Capital depriciatlon U Labour

H Fixed costs Earnings:

Heating oil Sugars Aromatics B Electricity (production) CB Steam

Pinewood

Fig. 13. Annual operating costs and earnings for the different cases at a feedstock capacity of 5 t h I For each individual case the first column shows the operating costs and the second column the earnings.

obtain any profit even for bagasse. However, the mismatch between the annual costs and the earnings is much less for the bagasse S&HOAL and S&P&HOAL cases compared to the HOu and HOAL cases.

5.3. Sensitivity study

A sensitivity study was carried out to identify the key factors influencing the economic feasibility. The parameter varied were the biomass price, sugar price, sugar yield, aromatic price, aromatic yield, heating oil price, hours of operation, the CAPEX, labor price and electricity price. The results are shown in Fig. 14. The project is economical feasible if the internal rate of return (IRR) is larger than the inflation and the interest or dividend paid to the project financiers. The IRR for the different bagasse cases at 50 t h_1 varied between 2.9 and 15.5%. The most important parameters for the production of only heating oil from untreated or acid leached

bagasse were the biomass price, heating oil price and the CAPEX. The labor price (largely depending on location) and hours of operation had a moderate impact. The parameter with major influence for the production of sugars together with heating oil or aromatics was the CAPEX. Reducing the CAPEX by 20% would increase the IRR to 21.4% and 22.6% for the production of sugars with heating oil and sugars with aromatics respectively.

6. Discussion and outlook

In this study pinewood, bagasse and straw were used as feedstocks with varying moisture, AAEM and extractives content. Obviously, the biomass price is one of the major cost drivers. For the S&HOAL cases at 50 t h-1 scale the biomass cost was 44% (pinewood), 35% (straw) and 28% (bagasse) of the total annual costs. This study showed that in case of acid leaching, processing biomass with

S.R.G. Oudenhoven et al. / Biomass and Bioenergy xxx (2016) 1—17

Electricity price (0.048; 0.06; 0.072 $ kWh| 1) Labour price (48000; 60000; 72000 $ year ) CAPEX (+20; 0;-20%)

Hours of operation (7000; 7500j 8000 h year "1) Heating oil (9.6; 12; 14.4 $ GJ" ) Aromatic yield on lignin (o.56; 0.7^0.84 kg kg ) Aromatic price (320; 400; 480 $ t" ) 1 Sugar yield on cellulose (0.3; 0^4; 0.5 kg kg ) Sugar price (280; 350; 420 $ t ) Biomass price (42; 35; 28 $ t )

Bagasse HOu at 50 th"1 I

Bagasse HO^ at 501 h"1 B

-10%-8% -6% -4% -2% 0% 2% 4% 6% 8% 10% Internal rate of return

-4% -2%

2% 4% 6% 8% 10% 12% 14% 16% Internal rate of return

Electricity price (0.048; 0.06; 0.072 $ kWI) 1) Labour price (48000; 60000; 72000 $ year" ) CAPEX (+20; 0;-20%)

Hours of operation (7000; 7500* 8000 h year_1) Heating oil (9.6; 12; 14.4 $ GJ" ) (

Aromatic yield on lignin (0.56; 0.7^0.84 kg kg ) Aromatic price (320; 400; 480 $ t" ) i

Sugar yield on cellulose (0.3; 0_4 0.5 kg kg ) Sugar price (280; 350; 420 $ t( ) Biomass price (42; 35; 28 $t )

Bagasse S&P&HO ^ at 501 h"1 I

4% 6% 8% 10% 12% 14% 16% 18% 20% 22% 24% Internal rate of return

8% 10% 12% 14% 16% 18% 20% 22% 24% 26% Internal rate of return

Fig. 14. Sensitivity study for the different cases for bagasse pyrolysis at 50 t h 1 scale.

a high moisture content does not lead to significant additional costs compared to dry feedstocks while. biomass with a high water extractives content will lead to large waste water treatment costs (straw 23% of total annual cost for S&HOAL case at 50 t h-1). Promising cheap biomass streams next to bagasse might be; empty fruit bunches, mallee wood and rice husk.

Using organic acids, produced by pyrolysis itself, instead of mineral acids for the acid leaching of the biomass has several advantages; i) no sulfur (in case of H2SO4) is added which will end up in the char, gas and/or oil ii) organic acid are expected to be less corrosive to the used equipment, especially at increased temperatures, iii) no hydrolysis reactions occur during leaching thus minimizing the organic loss to the waste water during the leaching process.

The energy balance showed that enough heat is produced from combustion of char and gases to provide the heat for pyrolysis and biomass drying. However, this only holds when the biomass is mechanically dried to approximately 50% by weight prior to thermal drying. This makes the dewatering press unequivocally one of the key operations in this concept.

Experiments performed on acid leaching of biomass gave a rough estimate of the leaching time, biomass moisture uptake, preferred biomass particle size and AAEM removal efficiency. Based on this information a diffuser (countercurrent moving bed extractor) was selected which is "readily" available. Further percolation tests and pilot testing, to optimize the process parameters, are required. A key factor in minimizing the water usage for the leaching process is the multi-stage counter-current operation mode. Note, that the AAEM content in the rinsing water must be sufficiently low, in our case 100 ppm, to effectively remove the AAEM from the biomass.

The waste water treatment cost in this work was based on a typical price for aerobic bacterial treatment (330 $ t-1 organic material). As mentioned in section 4.3, the COD content of the waste water streams indicates that implementing anaerobic digestion (CH4 production) before the aerobic bacterial treatment can reduce this cost. New developments, like fermentation to lipids [64] or microbial electrolysis to hydrogen [65] might increase the value of the organics in this stream even further. The removal of the inorganic, via e.g. reverse osmosis can reduce the fresh water

consumption of the plant. However, the cost of rinsing water is only 0.5% of the biomass feedstock cost (based on a process water cost of

0.2 $ m-3).

In our previous paper we showed that acid leached biomass has the tendency to melt during pyrolysis leading to char agglomerates in our fluidized bed reactor [72]. At 360 °C no operational problems caused by the melted biomass were observed, while still high organic oil and sugar yields were produced [72]. Since the pyrolysis reactions at 360 °C are relatively slow, the heating rate of the biomass particles is of less importance. Therefor reactors considered for slow or intermediate pyrolysis like augers and rotating kilns can be considered to produce high oil yields after acid leaching. However the slower reaction kinetics will require much longer residence times (min vs sec) of the biomass particles in the hot pyrolysis reactor which lead to increased CAPEX. The CAPEX and corresponding fixed costs (especially maintenance), were one of the major cost drivers of the process, e.g. for the S&HOAL cases at 501 h-1 scale the depreciation of the CAPEX plus fixed cost was 34% (pinewood), 29% (straw) and 52% (bagasse) of the total annual costs. This result shows the importance, and might challenge researchers and industry, to find cheap reactor concepts capable of dealing with melting biomass or long residence times. Moreover tests with existing reactor concepts, like the rotating cone reactor, auger reactor or circulating fluidized bed reactor, should be performed to study if the melting biomass leads to operational problems in these reactors types. A major advantage is that the produced char contains far less AAEMs which reduces problems (slagging and corrosion) in the combustor.

The heating oil (1st condenser oil or its fractions) produced from acid leached biomass will have different physical properties then normal pyrolysis oil (untreated whole oil), e.g. a low water content, less volatiles, less acidic and a much higher viscosity at room temperature [15]. This will require some adaption of used equipment like preheating of the oil before usage (~60 °C) to reduce the viscosity [15].

Currently, ethanol is still mainly produced by fermentation of simple sugars from predominately corn or sugar cane. Only recently a few lignocellulosic based processes have become commercial, which are based on pretreatment (e.g. acid hydrolysis, ammonia explosion) followed by enzymatic saccharification (using cellulase).

S.R.G. Oudenhoven et al. / Biomass and Bioenergy xxx (2016) 1—17

The advantages of this process compared to sugar production via pyrolysis are a higher cellulose to glucose yield (90% vs 40%) [6,68] and that a lower amount of inhibitors are formed. However the expensive cellulase (~0.1 $ dm-3 of ethanol) [68,73], long residence time (~3 days) [68], and relative low sugar concentrations (11% mass basis) [68] of the enzymatic saccharification are major drawbacks. The sugar content of the aqueous fraction after water and ethyl acetate extraction in our concept will be around 25—30% (see Fig. 8), which is higher than the sugar content obtained from enzymatic hydrolysis (10%—15% [68]). Moreover the direct fermentation of anhydrosugars (without hydrolysis to glucose) would avoid the use of mineral acids and prevent large amounts of solid waste from neutralizing. However, currently higher yields and working concentrations are obtained after hydrolysis of the anhy-drosugars [74]. It should be noticed that this comparison is rather obscured since the biocatalyst for anhydrosugars are hardly studied [57,74—77] whereas the biocatalyst for normal sugar fermentation are highly optimized. The IRR of 14.4% (S&HOAL) and 15.5% (S&P&HOAL), based on a sugar price of 350 $ ton-1, shows that the production of sugars via pyrolysis has clearly an economic potential. Another conversion method for anhydrosugars next to fermentation is acid-catalyzed conversion to produced levogluo-senone [78], furanics [79] or levulinic acid [79].

In case new insights, potentially gained from the currently well-studied lignin depolymerization, lead to improved techniques to produce monomeric phenolics from the (mainly oligomeric) phenolic fraction in pyrolysis oil would largely increase the economic value of this product slate and thereby the IRR of the whole process (a phenolic fraction price of 1000 $ t-1 gives an IRR of 30.5 for bagasse S&P&HOAL). Further research on the extraction of phenolic fraction from the 1st condenser oil is in this case advisable.

Finally, process integration with for example a sugar cane mill or palm oil mill can significantly reduce the CAPEX and OPEX (e.g. less operators are required). For example the CAPEX largely reduces when the char and gases can be fed to an existing combustor. Note, the char combustor is typical one of the most expensive devices in the pyrolysis plant (heat carrier loop ~50% of total CAPEX [22,80]). Another example of possible cost reduction due to integration is co-feeding of waste water in an existing (on-site) waste water treatment plant.

7. Conclusions

In this work the feasibility of removing AAEMs from biomass, using leaching with pyrolytic acids, as a pretreatment step in a pyrolysis process (5 and 50 t h-1 scale) has been evaluated in terms of mass and energy balance, economics and overall performance. A preliminary process design for the pyrolysis of pinewood, straw and bagasse is presented and an economic comparison was made between the pyrolysis of the untreated and acid leached biomasses.

Process simulation quantified that the pyrolysis with acid leaching can still be self-providing in its heat, required for the py-rolysis process and biomass drying, but insufficient excess heat is available to produce the electricity. The implementation of acid leaching will increase the total CAPEX of a new pyrolysis plant by 36% (only acid leaching) to 47% (including also product separation). For straw and bagasse an increase in the carbon to oil (straw 32% to 48%, bagasse 37%—56%) and energy to oil (straw 33%—51% bagasse 40%—57% of the initial biomass energy) was obtained after acid leaching but not for pinewood. The increased production of anhy-drosugars allows broadening of the product slate, which increases the product value (also for pinewood). under the current market circumstances and comparison of the three feedstocks, the production of heating oil and sugars from bagasse at a biomass scale of

50 t h-1 is the most economic option (IRR 15.4%).

It is not surprising that the economic feasibility largely depends on the biomass cost and CAPEX (process scale). For the acid leaching concept biomass with a high moisture content does not lead to a large cost penalty. Biomasses with a high content of water soluble extractives will lead to significant water treatment costs. Crucial in the leaching concept is the mechanical dewatering step which significantly reduces the water content of the biomass before thermal drying. Additionally, selecting a multi-stage counter-current extractor reduces significantly the rinsing water consumption.

The economics of the acid leaching concept can be further improved by: utilization of the phenolic fraction, process integration (e.g. combustor, waste water treatment) with existing industry like the sugar cane and palm oil mills and carefully selection of cheap biomass waste streams. To summarize, acid leaching of biomass combined with pyrolysis and fractional condensation offers opportunities for cost effective production of pyrolytic sugars, next to heating oil.

Acknowledgment

This research has been performed within the framework of the CatchBio program. The authors gratefully acknowledge the support of the Smart Mix Program of the Netherlands Ministry of Economic Affairs, Agriculture and Innovation and the Netherlands Ministry of Education, Culture and Science. The authors would like to thank Thomas Klein for his contribution to a part of the experimental work, Boudewijn de Smeth for the ICP analysis, De Smet for providing us with an estimate of the CAPEX and OPEX for the diffuser and the technical staff of the SPT group (Benno Knaken, Karst van Bree and Johan Agterhorst) for their excellent technical support.

Appendix A. Supplementary data

Supplementary data related to this article can be found at http:// dx.doi.org/10.1016/j.biombioe.2016.07.003.

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