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Liquefaction of lignocellulose in light cycle oil: A process concept study Shushil Kumar, Andrejs Segins, Jean-Paul Lange, Guus van Rossum, and Sascha R.A. Kersten ACS Sustainable Chem. Eng., Just Accepted Manuscript • DOI: 10.1021/ acssuschemeng.6b00055 • Publication Date (Web): 12 Apr 2016 Downloaded from http://pubs.acs.org on April 16, 2016

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ACS Sustainable Chemistry & Engineering

Liquefaction of lignocellulose in light cycle oil: A process concept study Shushil Kumara, Andrejs Seginsa, Jean-Paul Lange*a,b, Guus Van Rossuma,b, Sascha R.A. Kerstena a

Sustainable Process Technology, Faculty of Science and Technology, University of Twente, Drienerlolaan 5, 7522 NB, Enschede, The Netherlands

b

Shell Global Solutions International B.V., Shell Technology Centre Amsterdam, Grasweg 31, 1031 HW, Amsterdam, The Netherlands *Corresponding author: [email protected]; Tel: +31-20-630 3428

ABSTRACT Lignocellulosic bio-crude can be produced by direct liquefaction of lignocellulosic biomass, which can be further upgraded into biofuels in an oil refinery. Refinery streams, namely Vacuum gas oil (VGO) and Light cycle oil (LCO), were found suitable liquefaction solvents in our previous study. This paper reports a process concept based on the liquefaction of wood in LCO followed by recovery and recycling of the LCO. The LCO is recovered from the reactor effluent by spontaneous liquid-liquid phase split upon cooling. The feasibility of this process concept was demonstrated experimentally by conducting a series of refill experiments with LCO recovery and recycling. This series shows a steady-state liquefaction of pine wood

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ACS Paragon Plus Environment


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with constant product yields (liquid yield ~57 C%) and constant oil qualities over eight recycles. Most of the ash present in the wood was concentrated in the char. A technoeconomic assessment estimated that the bio-crude could be produced at 13.9 $/GJ, which corresponds to an energy-equivalent crude oil price of 61 $/bbl (at a wood cost of 85 $/t dry).

Keywords: Liquefaction, lignocellulose, biomass, light cycle oil, bio-crude, bio-oil, technoeconomic evaluation INTRODUCTION Lignocellulosic biomass is one of the promising renewable energy sources which can be converted into a liquid product (e.g. bio-crude) by using a number of different process routes13

. One of the process route to convert biomass into bio-crude is direct liquefaction which is a

variation of pyrolysis1, 4. Earlier direct liquefaction processes, namely those developed by the Pittsburgh Energy Research Center (PERC)4-5, and the Lawrence Berkeley Laboratory (LBL)6 as well as the Hydrothermal Upgrading (HTU)7-8 process, faced serious technical problems such as undissolved solids, increase of medium viscosity, poor oil quality etc.4. A few other processes were proposed more recently, namely STORS (Sludge-to-oil reactor system) process, CatLiq-process, and TCP (Thermal conversion process), by various research laboratories9-15. These processes were primarily developed to process a low-value wet organic feed stream e.g. agricultural and food waste, sewage sludge, manure etc., and, thereby, use water as a reaction medium. These processes might not be economical for a lignocellulosic feedstock (due to higher raw material cost), and will likely encounter similar issues as the HTU process. Our group16 revisited the direct liquefaction of lignocellulosic biomass using the produced bio-oil as a liquefaction medium (like PERC). Recycling the whole bio-oil initially succeeded

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in achieving a high oil yield (~90 C%) but readily lost its effectiveness as the liquid medium became very viscous due to build-up of heavy product (MW > 1000 Da) that slowly displaced the light start-up solvent16. Hence the follow-up work focused on chemical and physical approach to suppress build-up of the heavy product upon recycling through optimization of process parameters,17 the use of basic additives,18 or fractionation of the bio-oil into a light and a heavy fraction and recycle the lighter fraction19,20. In search for an alternative approach, we further explored refinery streams as an alternative liquefaction solvent, and found Vacuum gas oil (VGO) and Light cycle oil (LCO) as the two most promising solvents21. Both solvents produced a liquid yield of ~58% based on carbon and ~64% based on energy. We then proposed a very simple process that uses VGO, a typical FCC feed, in a once-through operation and sends the resulting VGO/bio-crude blend for further upgrading in an FCC unit21. This scheme was estimated to deliver bio-crude at 51-64 $/barrel of energy-equivalent crude oil. In the past a few attempts were made to explore refinery stream as a liquefaction medium by other researchers as well. The BioCRACK process22 liquefied biomass in VGO with a bio-oil yield of ~56 C% and a char yield of ~38 C% at 375OC. Jess et.al23 also studied the liquefaction of biomass in VGO and reported similar product yields. However, both studies provide limited information about the oil characteristics, the process conditions and eventual conversion or degradation of the liquefaction medium used.

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Figure 1: Process block diagram of the direct liquefaction of wood in LCO. The line thickness is indicative of the relative amount. Small amount of gas formed in the process is not shown for simplicity. This work builds up on the earlier findings of liquefaction in LCO21 (a typical FCC product stream), and demonstrates a process concept. The process concept is based on the liquefaction in LCO with subsequent recovery and recycling of the LCO as illustrated in Figure 1. The produced bio-crude was indeed found to separate out from the LCO upon cooling. The feasibility of this process concept is demonstrated by conducting a series of refill experiments with LCO recovery and its recycling. The effectiveness of recycle LCO as a liquefaction solvent is not known. Also the effect of the cross contamination of the solvent (with biocrude) and the bio-crude (with solvent) in the process needs to be investigated. In other words, is it possible to attain a steady-state process using the recycle LCO; has to be answered. Further, product characterization needs to be done in order to find a suitable use of the product (bio-crude) either as an end product or as a feed for further processing e.g. upgrading in a refinery into transportation fuels. Various analytical tools were used to characterize the product to monitor the approach to steady-state quality. Finally, a technoeconomic assessment of the process would help in assessing its economic potential. The resulting liquefaction process has the potential to be integrated within an oil refinery if product targets can be achieved economically. However, it can also be designed as a remote/distributed small-scale process located at a biomass production site; an option which

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was not possible with the process based on VGO proposed earlier21. A small amount of makeup LCO would still be needed to compensate a likely loss of LCO in the bio-crude. The produced bio-crude can then be transported to an oil refinery for further processing. A remote liquefaction process would reduce the transportation cost of biomass because of energy densification achieved upon converting biomass at 1.5-7.0 GJ/m3

24-25

to bio-crude (from

liquefaction) at 30-40 GJ/m3.10 It is contemplated that the bio-crude produced in the process can be co-processed with a fossil feed to obtain a product that can be blended in the refinery process chain. EXPERIMENTAL SECTION

Materials: Pine wood was obtained from Rettenmaier & Söhne GmbH (Germany). It was crushed to a particle size below 0.5 mm and then was dried at 105oC for 24 hours in an oven. The composition of the wood is listed in supporting information (SI). Light cycle oil (LCO) was obtained from Shell Technology Centre, Amsterdam, The Netherlands. Experimental set-up and procedure: Experiments were carried out in two different batch autoclaves having internal volumes of 45 mL and 560 mL. Safe operation was achieved by operating them in an explosion-proof high-pressure room with external control of the unit. The experimental set-ups of both reactors are described in detail elsewhere17. The major difference between the two autoclaves was their heating rates, which were around 75 and 3OC/min in the 45 mL and 560 mL autoclaves respectively. In this work, the 45 mL reactor was equipped with a 6 µm filter at the bottom of the reactor in order to filter the liquid from the solid at high temperature and high pressure (see schematic diagram in the SI). The detailed experimental procedure for the refill runs can be described as follows (a schematic diagram of experimental procedure can be found in the SI): The feed charge of 4 g of wood (pine) and 16 g of LCO (fresh LCO in first run) were loaded into the 45 mL reactor 5



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and mixed thoroughly. The reactor was closed tightly, flushed with nitrogen several times to remove any oxygen present in the system and pressurized to about 5 bar of nitrogen. The stirrer was turned on and the autoclave was immersed into a preheated fluidized sand bed which had a temperature set 50C above the desired reaction temperature. After the desired reaction time, the reactor was lifted out of the sand bath, the hot liquid and gas were passed through the filter and then through a cold water bath where vapors were condensed and cooled down. The liquid and incondensable gases were collected in two separate containers. The gas container was a special device similar to a U-tube manometer (filled with water saturated with KHCO3) with one end connected to the incoming gas line and the other end exposed to atmospheric pressure. The device allowed us to measure the total gas volume at a low pressure. After complete depressurization of the reactor, the autoclave was quenched in a cold water bath and subsequently cooled to ambient temperature. The remaining product mixture (solid soaked with liquid) was taken out from the reactor, washed with acetone and filtered with a 6µm filter. Acetone was then removed from the filtrate by vacuum evaporation, and from the solid residue by atmospheric drying at 105OC. The two liquids obtained after the hot filtration and from the acetone wash were mixed together. The resulting liquid split spontaneously in two phases, namely a large LCO-rich phase and a small bio-crude-rich phase. The bio-crude-rich phase was recovered as a reaction product. A small amount of the LCO-rich phase (normally 0.5 g) was taken out for analyses and the remaining part was used as a liquefaction solvent for the subsequent run. The LCO-rich phase was then added to the autoclave together with a second weighted amount of the dry pine wood (4 g) and a small amount of the make-up LCO to keep the total solvent amount of 16 g. The autoclave was then sealed and subjected to a second run. This procedure was repeated over 8 consecutive runs.

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The amount of wood and the total solvent were kept same (4 g wood and 16 g solvent) in all the runs. Experimental procedures of refill runs in the 560 mL reactor were same as in the 45 mL autoclave, except for the filtration which was performed outside the reactor at low temperature (70℃) and at low pressure (using 0.5 MPa of N2). The feed charge of 60 g of the wood and 240 g of the LCO (fresh LCO in first run) was used. An electrical jacket was used for heating and water was sprayed on the outer reactor wall to cool the reactor after the reaction. The reaction temperature was defined as the end temperature and the reaction time was defined as the time the autoclave spent in the hot sand bath (including heating time of ~5 min or ~60 min for the 45 mL or 560 mL autoclaves, respectively). Analyses: Gas samples were analyzed with an off-line gas chromatograph (Varian Micro GC CP-4900). The liquid product was analyzed with a Gel Permeation Chromatograph (GPC; Agilent 1200 series, with RI and UV (wavelength: 254 nm) detectors). More details about the equipment can be found elsewhere18. The elemental composition was determined using Elemental Analyzer (Interscience Flash 2000). The chemical nature (type of bonds) was investigated by means of Fourier Transform Infrared Spectrophotometer (FT-IR Bruker Tensor 27). pH and TAN (total acid number) were measured using an autotitrator (785 DMP Titrino, Metrohm). MCRT analysis was performed following the ASTM D4530 standard. The minerals in char, bio-oil and pine wood were analyzed by using portable XRF (Niton CL3t GOLDD+) with measuring time of 30 seconds for main, low and light filters, and 5 seconds for high filter. For the analysis of minerals, the sample was first calcined at 600OC in air to obtain ash which was used for the analysis. The viscosity was determined by Viscosity meter (Brookfield DV-E Viscometer) and the water content was determined by Karl-Fischer

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titration (Metrohm 787 KF Titrino, with Hydranal composite as titrant). The higher heating value (HHV) was measured using a Bomb calorimeter (IKA C2000 basic). Product definition and calculation: Yields were defined as C% of the biomass intake using common definitions explained elsewhere17. In short, the gas and solid yields were determined directly via GC analysis of the off gases and by weighting and elemental analysis of the solid after filtering, washing and drying. In contrast, the liquid yield is determined indirectly as ‘100 - gas yield – solid yield’, as direct measurement was not possible due to cross contamination of the LCO and the bio-crude. The liquid products (LCO-rich phase and bio-crude-rich phase) consisted of two fractions that were defined based on apparent molecular weight (as determined by GPC), namely solvent (LCO) with MW,GPC < 150 Da and the bio-crude with MW,GPC > 150 Da. The bio-crude fractions of each phase contained ‘Vacuum residue’ (VR) defined as MW,GPC >1000 Da. The vacuum residue fraction was defined as the fraction of the bio-crude-rich phase that is found in the vacuum residue (MW,GPC >1000 Da) based on equation 1. It assumes comparable response factor for the light and heavy bio-crude components. =

, !" #$%%% & '( ) *+, -.

(1)

It should be mentioned that the apparent molecular weight of the vacuum residue corresponds to that of refinery vacuum residue16. Refinery terminology is used here to facilitate the translation into the refinery operation to be selected for upgrading the bio-crude into final biofuels. The contamination of the LCO-rich phase with the bio-crude, i.e. its ‘Bio-crude fraction’, was defined as the fraction of components with MW,GPC > 150 Da in the LCO-rich phase (equation 2).

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/ − =

, !" #$1% & '( ) *+, -.

(2)

A similar contamination of the bio-crude with LCO was evaluated as a ‘LCO fraction’, and was defined as the fraction of components with MW,GPC < 150 Da in the bio-crude-rich phase (equation 3).

234 =

, !" 5$1% & '( ) *+, -.

(3)

RESULTS LCO was found to be one of the promising liquefaction solvents, after screening various refinery streams in our previous study21. It also offered the convenience of spontaneously separating out from the bio-crude-rich phase upon cooling. The optimization of process conditions for the liquefaction (see Kumar et.al.21) led to selection of the following process conditions for this study. 1. Temperature ~320℃ 2. Reaction time ~15 min 3. Solvent:Wood = 80:20 wt%, water addition was not required Feasibility of the process concept (Figure 1) was explored by conducting refill experiments in the LCO-rich phase isolated in a preceding run. Refill experiments in the batch autoclaves were intended to mimic a continuous operation. The focus of the refill experiments was to approach a steady-state operation in terms of product yields and product characteristics. Cross-contamination of the LCO (with the bio-crude) and the bio-crude (with LCO) and their effects on the liquefaction efficiency were also investigated.

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Product yields: The product yields of the refill experiments carried out in the 45 mL autoclave are shown in Figure 2. The yields appear to be similar in all the refill runs. The liquid yield was around 57 C% while solid yield was around 40 C%. Gas yields were very low at around 3-5 C% and consisted mainly of CO2 (~60 vol%) and CO (~40 vol%). The use of a larger autoclave (560 mL) led to slightly different yields; the liquid and solid yields were around 51 and 44 C% respectively (see the SI). The lower liquid yield can be attributed to the slower heating and cooling rate of the larger autoclave, which resulted in a longer reaction time (100 min vs. 15 min in the hot sand bath) and favored secondary reactions to char formation. The energy content of the wood (based on Higher heating value (HHV)) was distributed for ~62 % to the liquid, ~35 % to the char and 1 % to the gas, after correcting for ~2 % of the energy released as heat of reaction (as found in the earlier study with VGO21, see the SI for detail). The water content in the bio-crude-rich phase was measured to be 7-9 wt%, while the LCO-rich phase was free of water, within the detection sensitivity. In summary, a stable liquefaction process in terms of products distribution can be achieved, when using the recovered LCO-rich phase as a liquefaction medium.

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60 Liquid

50

Yield (C%)

40

Solid

30

20

10 Gas

0

St ar t-u p 1s tr ef ill 2n d re fil l 3r d re fil l 4t h re fil l 5t h re fil l 6t h re fil l 7t h re fil l 8t h re fil l

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 2: Product yields obtained in the refill runs using the separated LCO-rich phase in 45 mL autoclave, τ = 15 min. Feed (wt%); Solvent:Wood = 80:20, T~320OC. In this work replication of experimental results is not carried out. However, the refill experiments imply replication. After the few initial refill experiments, the remaining refill experiments can be regarded as replication, as nothing was changed. The major changes in the recycle LCO could be observed only during the first three refill experiments (see Figure 3), and after that it reached steady state. The amount of feed components and process conditions were kept the same in all the refill experiments. The results of refill experiments show good reproducibility of the experimental results. Product characterization: The liquid product separated in two liquid phases, which were characterized in order to assess their qualities. The bottom phase consisted mainly of high molecular weight components, while the top phase consisted mainly LCO (see the SI for GPC traces).

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Vacuum residue fraction, contamination and chemical functionality: The quality of the bio-crude-rich phase appeared not to change upon successive runs with the recycle LCO. Figure 3 (top) shows no significant change in the vacuum residue (VR) and LCO fractions with number of refill runs in both the autoclaves. A slight dip in the VR fraction and a rise in the LCO fraction in the 6th refill were possibly due to one month delay in the analysis of the sample. This likely led to deposition of the heaviest fraction of the bio-crude-rich phase on the wall of storage vial, making the bio-crude-rich phase richer in the light components and leaner in the heavy components.

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0,8

VR fraction

0,3

0,6

0,2 0,4

0,1

0,2 0,0

0,0 Bio-crude fraction

LCO fraction

1,0

0,4

0,3 0,2 Pure LCO

0,1 0,0

C=O/C-H

0,3 0,2 0,1

re fil l 4t h re fil l 5t h re fil l 6t h re fil l 7t h re fil l 8t h re fil l

re fil l

3r d

2n d

up

1s t

re fi l l

Pure LCO

0,0

St ar t-

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 3: Vacuum residue and LCO fractions in the bio-crude-rich phase (top), bio-crude fraction in the LCO-rich phase (middle) and relative carbonyl content of the LCO-rich phase based on ratio ‘C=O band (1709 cm-1) / C-H band (2900 cm-1)’ in FT-IR spectra (bottom) with number of refill runs. Open symbol shows results in the 560 ml autoclave while solid symbol shows results in the 45 mL autoclave. Solvent:Wood = 80:20, T = 320℃, τ = 15 min for 45 mL and 100 min for 560 mL. The LCO-rich phase stream appeared to remain of constant quality as well (see the SI for GPC figures). Its bio-crude fraction increased initially but stabilized after the 3rd refill (Figure

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3 - middle), which indicates saturation of the LCO-rich phase with the bio-crude within three refill runs. It was ensured that the trend was independent of the choice of molar mass cutoff for the bio-crude fraction. Furthermore, FT-IR analysis showed a rapid and modest saturation of the carbonyl content of the LCO-rich phase, as measured by the relative intensities of C=O band (1709 cm-1), which is a characteristic for the bio-crude components (aldehydes, ketones acids and/or esters), to C-H band (2900 cm-1), which is a characteristic for the LCO components (Figure 3 - bottom). The characteristics of the bio-crude-rich phase and the LCO-rich phase obtained in both autoclaves show similar trends but differ slightly. Compared to the smaller autoclave, the larger autoclave produced a bio-crude-rich phase with a higher VR fraction and an LCO-rich phase with a higher bio-crude contamination as apparent by its higher bio-crude fraction and C=O/C-H ratio. All these effects may tentatively be attributed to the higher reaction time of the larger autoclave and enhanced secondary reactions. Elemental composition and heating value: Elemental analysis and heating value of the feed and the product streams after the 8th refill run are shown in Table 1. A significant reduction of 68 % in oxygen content has been observed by converting the wood into the biocrude-rich phase. Also the bio-crude-rich phase (dry) has a 56 % higher energy content than the wood. A significant part of these changes is likely coming from the ~20 % LCO present in it (Figure 3). The char has a higher energy density than the feed wood as well and, hence, can be used as a process fuel. The bio-crude-rich phase has a low H/Ceff of 0.7 (Table 1) which suggests a high coke yield when processing in an FCC unit. For instance, feedstocks with H/Ceff ratios below 1 are claimed to be difficult to upgrade to premium products over a ZSM-5 catalyst due to rapid aging and deactivation of the catalyst26. The bio-crude-rich phase is not likely to be suitable as

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a stand-alone feed. It will rather need to be mixed with other refinery stream to uplift the H/Ceff ratio. Petroleum-derived feedstocks typically have H/Ceff ratio between 1 and 2. Table 1: Elemental compositions and heating values of different process streams obtained in the 8th refill run in the 45 mL autoclave. Elemental composition (dry wt%)

H/Ceffb

Heating valuec

C

H

Oa

Wood

46.6

6.3

47.1

0.11

19.2

Bio-crude-rich phased

78.7

6.4

14.9

0.70

29.9

Char

68.8

5.1

26.1

0.32

26.6

LCO-rich phasee

85.8

10.4

3.2

1.40

42.3

(dry, MJ/kg)

a

by difference, bH/Ceff = (H-2×O)/C; here H, O and C are atomic composition, cmeasured using a bomb calorimeter, dpartly affected by the LCO present (~20 wt%) in the bio-crude-rich phase, eelemental composition of pure LCO: C:H=86.3:10.3.

Viscosity, TAN, MCRT, Ash and miscibility in VGO (Vacuum gas oil): The small amount of samples obtained in the refill experiments in the 45 mL autoclave did not allow further analyses to be carried out. Therefore, further analyses of the liquid product were limited to the samples which were obtained in the larger autoclave, and shown in Table 2. The viscosity of the bio-crude-rich phase was found at least ~17 times higher than the viscosity of a pyrolysis oil (with 5 % water content ) of ~30 cP measured at 70℃27. The viscosity of the LCO-rich phase did not change with refill runs and remained very low at ~4 cP (25℃). The acidity of reactants and products is an important factor in an industrial design, as it determines the corrosiveness and, therefore, the choice of material for the equipment. The acidity of the bio-crude-rich phase did not change significantly with the refill runs and remained around 50-60 mg/g, slightly below the TAN of pyrolysis oil at around 7028-29. The

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TAN of the bio-crude-rich phase indicates a mild acidity, which implies that stainless steel (type 316) could be a suitable material of construction for the process equipment handling the bio-crude-rich phase30. The TAN of the LCO-rich phase in all the refill runs was very low at < 10 mg/g. The TAN was measured at neutralization pH of around 8.5. Table 2: Characteristics of different process streams obtained in the 4th refill run in the 560 mL autoclave. Bio-crude rich phase

LCO rich phase

Wood

Char

500 at 70OC

4 at 25OC

-

-

TAN (mgKOH/g)

60

10

-

-

MCRT (wt%)

26

N/A

-

-

0.18

0.02

0.5

3

1

-

9

23

Viscosity (cPa)

Ash contents (wt%) TBN (mgKOH/g) a

measured at spindle speed of 40 revolution per minute. A micro carbon residue test (MCRT) was conducted to measure the coking tendency of the

bio-crude-rich phase. The MCRT value of different bio-crude-rich phases obtained during the refill runs varied between 25 and 33 wt% (see the SI). The MCRT of the bio-crude-rich phase was comparable to the 26 wt% MCRT of pyrolysis oil31. It should be noted here that such values are significantly higher than the 0.2 wt% of an FCC feedstock (vacuum gas oil; VGO)32. Therefore, further processing of the bio-crude-rich phase in an FCC unit may require to mix the bio-crude-rich phase with a large portion of VGO in order to avoid/minimize issues that may arise due to large coke formation from the bio-crude. The MCRT of the mixture of

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bio-oil and VGO was reported much lower than the linear combination of the MCRT of the two individual components33-34. Wood contains a small amount of ash which consists of different types of minerals. Proper understanding of the fate of minerals is critical for two reasons. Firstly, the minerals are essential for the plant growth and therefore, should be recovered for recycling into the soil as much as possible. Secondly, basic constituents that would make their way into the bio-cruderich phase may affect its further processing, e.g. by neutralizing the acid sites of the FCC catalyst. The measurements of ash contents in various streams show that the ash was mainly concentrated in the char with a concentration of ~3-4 wt% vs. only 0.18 wt% in the bio-cruderich phase and 0.02 wt% in the LCO-rich phase. The detailed analysis of the ash further allows us to quantify the basicity or neutralization power of the bio-crude-rich phase (see the SI). Accordingly, the concentration of basic components Na, K, Ca, Mg and Mn in the wood, bio-crude-rich phase and char were converted to KOH equivalent and were summed up to a total base number (TBN), with the assumption that 1 mole of divalent cation corresponds to 2 moles of KOH. The TBN of the wood, bio-crude-rich phase and char were found to be 9, 1 and 23 mg of KOH/g of sample respectively. Interestingly, the bio-crude-rich phase shows a very low TBN (1 mgKOH/g) and will therefore have a limited poisoning effect on the FCC catalyst. Finally the miscibility of the bio-crude-rich phase in the VGO (a FCC feed) was investigated by placing the bio-crude-rich phase and the VGO in a quartz capillary and then heating it in an oven. Mixing was done by shaking the capillary. A complete homogeneous solution was not observed until 465℃ which is well above the feed temperature of FCC unit (350-390℃) (see SI). However, a significant part of the bio-crude-rich phase appeared to dissolve in the VGO at 255℃. The high oxygen content (~20 wt%) and very high molecular

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weight of the bio-crude-rich phase are likely responsible for its poor solubility in VGO even at that high temperature. Therefore, the likely option would be to homogenize the bio-cruderich phase and VGO prior to feeding them to the FCC reactor. PROCESS SIMULATION AND TECHNO-ECONOMIC ANALYSIS At these very early stage of the research, it was worth considering to gain preliminary views on the economic viability of the process, and hence to identify the aspects of the process that are most critical for the economic viability. Therefore a very preliminary approach, mainly based on product yields, and energy transfer duty as a first approximate for inside battery limit (ISBL) cost was considered in this study. Process simulation and energy requirements: The economic potential of the process concept has been evaluated by determining its energy demand and energy transfer requirements through process modeling with Aspen Hysys. In the model, the liquefaction reactor was fed with solvent and wood in 80:20 weight ratio, as applied in the experimental studies. The resulting product yields were set at 68 wt% liquid, 6 wt% gas and 26 wt% solid as found in the recycle experiments (see the SI). The water yield was set at 20 wt% (based on wood) as found in the earlier VGO process21. This resulted in an organic yield of 48 wt%. It is likely that a small amount of organics will be present in the aqueous phase as observed in our earlier work17. The process simulation used “Basra Light” petroleum assay from the Aspen Hysys library. A distillation cut of 120-320OC from the “Basra light” was used for the LCO stream. Wood and the bio-oil were also simulated as the LCO. The process flow diagram can be found in the supporting information. Lignocellulosic feed is liquefied in a liquefaction reactor. The reactor effluent was first sent to a filtration unit to

18


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filter the solid before passing through a heat exchanger to be cooled while heating the feed mixture. Thereafter the resulting cold reactor effluent (at ~70℃) was sent to a decanter to separate the LCO-rich phase from the bio-crude rich phase. The temperature of ~70℃ was assumed to be sufficient for the phase split. The filtered solid will entrap some liquid, which can be recovered by vacuum drying but was modelled here using a flash column at atmospheric pressure. The char was assumed to contain twice its weight of liquid after filtration but, contained only 20 wt% of its weight after vacuum drying. This results in a liquid yield loss of 5 wt% (based on wood). The energy requirement of the process was calculated to be amounting to ~3 % of HHV of the wood intake, with a breakdown of 1 % for heating the feed and 2 % for drying the solid. Such energy demand can easily be met by the by-product char, which accounts for ~35 % of the energy fed as the wood. The total heat transfer amounts to 4 GJ/ton of the wood. Techno-economic assessment of the liquefaction process concept: The capital cost (inside battery limit cost only (ISBL)) of the process could be estimated from the energy transfer duty of the individual process equipment calculated above, using the correlation suggested by Lange35 as given by equation 4. The correlation is based on the ISBL cost of existing technologies. Contingency for new technology is not considered at this stage. Equation 4 is updated by replacing the original factor of 2.9 with 4.7 to account for cost escalation from year 1993 to 2014*. The capacity of the plant was arbitrarily set to 100 t/h of dry wood. 67 869/2, :$ 2014@ = 4.7 × D EF

8:G@H%.11

(4)

*

using Chemical engineering plant cost index; CEPCI for year 2014 (July) is 576.9 and for year 1993 is 359.2.

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The ISBL cost of the process was estimated at ~100 M$. The total investment cost of the process was obtained by multiplying the ISBL cost with a factor of 2.26 that takes other costs into account (detail calculation can be found elsewhere20). The overall production cost is shown in Table 3 (for detail calculation see Kumar et.al20). The liquid lost on the char (5 wt% based on the wood intake) was assumed to come from the bio-crude and hence subtracted from the total yield of the bio-crude (this assumes similar prices of LCO and the bio-crude). The energy requirement of the process was met by burning the char produced during the process. The remaining char (along with trapped liquid) was credited for its energy value at a price of 50 $/t of the char; a price similar to coal. The wood was priced at 85 $/t (including transportation cost), as used in other studies on biomass processes36-37. Table 3: Process economics data for the direct liquefaction of biomass in LCO, with recovery and recycling of the LCO. Operation time: 8320 h/year. Feed intake

kt/year of wood (dry) t/h

832 100

Product (bio-crude)

t/h

43a

Bio-crude yieldb

t/t

0.43

C/C

0.53

GJ/GJ

0.56

Feed price

$/t (dry)

85

$/GJ

4.6

Energy credit

M$/year

-12

Total investment

M$

222

Capital charge

M$/year

26

Total production cost

M$/year

123

$/t of bio-crude

349

$/GJ of bio-crude

13.9

c

Break-even crude price a

$/bbl

61 b

5% liquid lost on the char is assumed to come from the bio-crude, based on the wood, cEnergy credit in form of char produced at 50 $/ton of the char (similar to coal price).

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The total investment cost was estimated to be 222 M$, which is comparable to the ~210 M$ calculated for a pyrolysis process and lower than the ~360 M$ calculated for a liquefaction process (after rescaling to 100 t/h and updating to year 2014)38. Other studies also reported significantly higher investment costs for various cellulosic biofuel plants, namely 360M$39-380M$40 for a pyrolysis plant with subsequent upgrading of the bio-oil, 300500M$36, 41-43 for a cellulosic bio-ethanol plant and 500-600M$36 for a gasification/FischerTropsch plant, all after rescaling to 100 t/h and update to year 2014. It is fair to stress that the 222 M$ investment estimated here don't consider the cost (and yield loss) of upgrading the bio-crude to final fuel. The production cost of the bio-crude was estimated to be 349 $/t or 13.9 $/GJ, which corresponds to an energy equivalent crude-oil price of 61 $/bbl (Table 3). The production cost can be divided into three components that are (i) variable operating cost (feedstock, utility) of 48%, (ii) fixed operating cost (labor, maintenance, overheads, taxes, insurance, lab, royalty) of 31% and (iii) capital charge of 21%. The capital charge was estimated with the annuity method employing a 10% interest rate and a service life of 20 years as used by Elliott et.al.44 for the techno-economic assessment of biomass liquefaction processes. A similar breakup of the three components was also found in their studies. The production costs for a fast pyrolysis process and for a liquefaction process reported by Elliott et.al.38 were 10.5$/GJ and 15.5$/GJ (wood price: 60 $/dry ton, currency year 1991) using similar methods. Similar production costs of 13.8 $/GJ (61 $/bbl of crude oil equivalent) and 14.5 $/GJ (64 $/bbl of crude oil equivalent) were estimated for liquefaction of wood in light bio-oil20 and in VGO21, respectively.

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A further reduction in the investment cost and hence in the production cost of the bio-crude can be achieved by omitting the washing step of char as demonstrated in the previous study based on the VGO21. The production cost depends mainly on the feed price, bio-crude yields, investment cost and by-product credit. For our sensitivity analysis, we considered a relative variation of ±30 % for feed and by-product prices, ±50 % for capex (capital expenditure/capital charge) and ±15 % for the bio-crude yield. Consequently, the bio-crude production cost could vary within a band of ±13 $/bbl of oil equivalent (Figure 4). Break-even crude oil price ($/bbl) 45

50

55

60

65

Feed price (85 $/t)

70

75

±30%

Bio-crude yield ( 57 C%)

±15%

Capex (220 M$)

±50%

By-product (char) credit (50 $/t)

±30%

+ -

Figure 4: Sensitivity analysis of production cost of liquefaction bio-crude, expressed in breakeven crude oil price. The base case values are reported between brackets; the sensitivities are expressed as relative percentages. The production of the bio-crude from this process is competitive if not better than other processes to produce bio-fuels. The cost estimation of the liquefaction process here is very preliminary. However, given the promising results of preliminary techno-economic assessment, it is worth further study.

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CONCLUSION This work demonstrates a process concept based on the liquefaction of wood in LCO with recovery and recycling of the LCO. From this work the following conclusions can be drawn: • A steady-state liquefaction of pine wood with constant product yields and oil qualities over several recycles was demonstrated. • A stable concentration of the bio-crude in the recycle LCO stream was observed after a few refill runs. • The bio-crude-rich phase is significantly heavy, acidic (TAN ~60), highly viscous (500 cP at 70℃), has high coking tendency (~26 wt% MCRT) and a moderate ash content around 0.2 wt%. • Ash present in the wood was mainly found in the char, which makes easier to recycle the minerals back to soil either by burning the char or by using the char as a fertilizer. • The techno-economic assessment of the liquefaction process suggests that the bio-

crude could be produced at 12.4-15.8 $/GJ, which corresponds to an energyequivalent crude oil price of 55-70 $/bbl. The production cost shows the highest sensitivity to the feedstock price and the bio-crude yield. Investment cost seems much less critical at this stage. AUTHOR INFORMATION Corresponding Author *E-mail: [email protected] Notes The authors declare no competing financial interest. ACKNOWLEDGEMENTS

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The authors would like to thank Shell Global Solutions International B.V. for funding this research and Benno Knaken for the technical support. SUPPORTING INFORMATION Wood characterization (Table S1); experimental setup and procedure (Figure S2-S3); Product yields based on wt% (Figure S3) obtained in 45 mL autoclave; Product yields based on C% (Figure S4) obtained in 560 mL autoclave; GPC graphs of various liquid streams (Figure S5); Measured viscosity (Figure S6), TAN (Figure S7), MCRT (Figure S8) and Ash contents (Figure S9) of various components; Minerals concentration in the ashes (Table S2); Miscibility of the bio-crude-rich phase in VGO at different temperatures (Figure S10); Process flow sheet of the process concept (Figure S11); Material balance around the reactor (Table S3); Heat duty of heat transfer equipment (Table S4). This material is available free of charge via the Internet at http://pubs.acs.org. ABBREVIATONS Da

Dalton

GPC

Gel Permeable Chromatography

HHV

Higher heating value

MW,GPC

Molecular weight defined by GPC

MW

Molecular weight

VR

Vacuum residue

SI

Supporting information

t

Metric ton

T

Temperature

τ

Reaction time

bbl

Barrel

FCC

Fluid catalytic cracker

LCO

Light cycle oil

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VGO

Vacuum gas oil

TAN

Total acid number

TBN

Total base number

MCRT

Micro carbon residue tester

XRF

X-ray fluorescence

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(32) Lappas, A. A.; Bezergianni, S.; Vasalos, I. A. Production of biofuels via co-processing in conventional refining processes. Catalysis Today 2009, 145 (1–2), 55-62. (33) de Miguel Mercader, F.; Groeneveld, M. J.; Kersten, S. R. A.; Way, N. W. J.; Schaverien, C. J.; Hogendoorn, J. A. Production of advanced biofuels: Co-processing of upgraded pyrolysis oil in standard refinery units. Applied Catalysis B: Environmental 2010, 96 (1-2), 57-66. (34) Samolada, M. C.; Baldauf, W.; Vasalos, I. A. Production of a bio-gasoline by upgrading biomass flash pyrolysis liquids via hydrogen processing and catalytic cracking. Fuel 1998, 77 (14), 1667-1675. (35) Lange, J.-P. Fuels and Chemicals Manufacturing; Guidelines for Understanding and Minimizing the Production Costs. CATTECH 2001, 5 (2), 82-95. (36) Anex, R. P.; Aden, A.; Kazi, F. K.; Fortman, J.; Swanson, R. M.; Wright, M. M.; Satrio, J. A.; Brown, R. C.; Daugaard, D. E.; Platon, A.; Kothandaraman, G.; Hsu, D. D.; Dutta, A. Techno-economic comparison of biomass-to-transportation fuels via pyrolysis, gasification, and biochemical pathways. Fuel 2010, 89, Supplement 1 (0), S29-S35. (37) Jones, S.; Male, J. Production of Gasoline and Diesel from Biomass via Fast Pyrolysis, Hydrotreating and Hydrocracking: 2011 State of Technology and Projections to 2017; Pacific Northwest National Laboratory: 2012. (38) Solantausta, Y.; Beckman, D.; Bridgwater, A. V.; Diebold, J. P.; Elliott, D. C. Assessment of liquefaction and pyrolysis systems. Biomass and Bioenergy 1992, 2 (1–6), 279-297. (39) Wright, M. M.; Daugaard, D. E.; Satrio, J. A.; Brown, R. C. Techno-economic analysis of biomass fast pyrolysis to transportation fuels. Fuel 2010, 89, Supplement 1 (0), S2-S10. (40) SB, J.; Valkenburt, C.; Walton, C.; Elliott, D.; Holladay, J.; Stevens, D.; Kinchin, C.; Czernik, S. Production of Gasoline and Diesel from Biomass via Fast Pyrolysis, Hydrotreating and Hydrocracking: A Design Case; PNNL-18284, Pacific Northwest National Laboratory, Richland, WA: 2009. (41) Humbird, D.; Aden, A. Biochemical Production of Ethanol from Corn Stover: 2008 State of Technology Model; 2009. (42) Humbird, D.; Davis, R.; Tao, L.; Kinchin, C.; Hsu, D.; A, A.; Schoen, P.; Lukas, J.; Olthof, B.; Worley, M.; Sexton, D.; Dudgeon, D. Process Design and Economics for Biochemical Conversion of Lignocellulosic Biomass to Ethanol : Dilute-Acid Pretreatment and Enzymatic Hydrolysis of Corn Stover; National Renewable Energy Laboratory Golden, Colorado: 2011. (43) Kazi, F. K.; Fortman, J. A.; Anex, R. P.; Hsu, D. D.; Aden, A.; Dutta, A.; Kothandaraman, G. Techno-economic comparison of process technologies for biochemical ethanol production from corn stover. Fuel 2010, 89, Supplement 1 (0), S20-S28. (44) Elliott, D. C.; Baker, E. G.; Beckman, D.; Solantausta, Y.; Tolenhiemo, V.; Gevert, S. B.; Hörnell, C.; Östman, A.; Kjellström, B. Technoeconomic assessment of direct biomass liquefaction to transportation fuels. Biomass 1990, 22 (1–4), 251-269.

For Table of Contents Use Only

Liquefaction of lignocellulosic in light cycle oil: A process concept study

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Shushil Kumar, Andrejs Segins, Jean-Paul Lange, Guus V. Rossum, and Sascha R.A. Kersten Synopsis: A direct liquefaction process converts wood to bio-crude in a refinery stream at a high yield (~60 C%). The resulting bio-crude has low oxygen content and high energy density.

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Liquefaction of Lignocellulose in Light Cycle Oil: A ... - ACS Publications

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