Extraction and purification of renewable chemicals from hydrothermal

Subscriber access provided by UNIV OF TEXAS DALLAS

Process Systems Engineering

Extraction and purification of renewable chemicals from hydrothermal liquefaction bio-oil using supercritical carbon dioxide: A techno-economic evaluation wahab maqbool, Kameron Dunn, William O. S. Doherty, Neil McKenzie, Dylan John Cronin, and Philip Hobson Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b05366 • Publication Date (Web): 07 Mar 2019 Downloaded from http://pubs.acs.org on March 8, 2019

Just Accepted “Just Accepted” manuscripts have been peer-reviewed and accepted for publication. They are posted online prior to technical editing, formatting for publication and author proofing. The American Chemical Society provides “Just Accepted” as a service to the research community to expedite the dissemination of scientific material as soon as possible after acceptance. “Just Accepted” manuscripts appear in full in PDF format accompanied by an HTML abstract. “Just Accepted” manuscripts have been fully peer reviewed, but should not be considered the official version of record. They are citable by the Digital Object Identifier (DOI®). “Just Accepted” is an optional service offered to authors. Therefore, the “Just Accepted” Web site may not include all articles that will be published in the journal. After a manuscript is technically edited and formatted, it will be removed from the “Just Accepted” Web site and published as an ASAP article. Note that technical editing may introduce minor changes to the manuscript text and/or graphics which could affect content, and all legal disclaimers and ethical guidelines that apply to the journal pertain. ACS cannot be held responsible for errors or consequences arising from the use of information contained in these “Just Accepted” manuscripts.

is published by the American Chemical Society. 1155 Sixteenth Street N.W., Washington, DC 20036 Published by American Chemical Society. Copyright © American Chemical Society. However, no copyright claim is made to original U.S. Government works, or works produced by employees of any Commonwealth realm Crown government in the course of their duties.

Page 1 of 28 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

Industrial & Engineering Chemistry Research

Extraction and purification of renewable chemicals from hydrothermal liquefaction bio-oil using supercritical carbon dioxide: A techno-economic evaluation Wahab Maqbool, Kameron Dunn, William Doherty, Neil Mckenzie, Dylan Cronin, Philip Hobson* Queensland University of Technology (QUT), 2 George Street, Gardens Point, 4000 Brisbane, Australia Correspondence E-mail: [email protected]

Abstract Supercritical fluid extraction (SFE) and fractionation of products from a complex mixture such as bio-oil, where many compounds are present in low concentrations, is a difficult process to model. This difficulty arises from the uncertainty associated with those interactions between mixture components for which fundamental vapour-liquid equilibrium (VLE) data is not available. In this work a novel extraction and purification concept is investigated using a predictive model developed from VLE data of binary solute-solvent systems; solute-solute interactions in the supercritical carbon dioxide (scCO2) phase are neglected. The predictive component of the work employs an equation of state (EOS) model to achieve the above task. The results of pilot plant trials utilising a bio-crude feedstock were shown to be in good agreement with the model predictions. Aspen Plus® process simulations were developed for the extraction process which comprised of supercritical extraction and subsequent purification steps utilising distillation and multistage evaporation. A techno-economic analysis of different process designs were evaluated for comparison. In particular, distillation as the primary separation process with and without multistage evaporation were simulated to compare the economics of supercritical extraction to distillation. It was found from simulation results that distillation is a very energy intensive process, and total operating costs for it are always greater than supercritical extraction counterparts. Combining multistage evaporation with distillation will bring the total operating cost slightly lower than supercritical extraction processes. However, SFE will extract products with better efficiency, yielding a much better annual total profits than distillation. Even using solvent/bio-oil (S/B) ratio of up to 20 will not have considerable impact on total profits of SFE process in relation to distillation.

1. Introduction Supercritical fluid extraction is currently in use for a number of niche applications1, 2 such as the decaffeination of coffee or the recovery of essential oils and bioactive compounds from plant materials. The use of SFE for the extraction of compounds from bio-oil has been the subject of a limited number experimental studies3-13. The lack of fundamental

ACS Paragon Plus Environment

Industrial & Engineering Chemistry Research 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

investigations into SFE of bio-oil can be attributed to the highly complex nature of bio-oil and the difficulty this presents in describing this process in terms of phase equilibria. Biooils are made up of large portions of water and many other chemical compounds but the latter only in small quantities.4, 6, 7, 13 A fundamental modelling approach based on an equation of state was adopted in the current study to investigate the novel SFE and subsequent purification of bio-oil compounds. The developed model for multicomponent mixture was used to determine the subsequent staged depressurization conditions required for the recovery of individual compounds or groups of compounds from the supercritical extract phase. In this work, in-house produced bio-oil from hydrothermal liquefaction (HTL) of black liquor, also known more commonly as bio-crude, was first extracted with scCO2 and subsequently fractionated in to two product fractions with the use of stage-wise pressure reduction techniques. In the currently proposed extraction process the highly dilute bio-crude in water feedstock is first fed in to the base of a SFE extraction column. Literature review and preliminary experiments helped to determine the conditions of temperature, pressure and bio-crude pH at which the SFE of our bio-crude from the aqueous phase will produce equilibrated extract samples in pilot plant trials. The supercritical extract stream emerging from the top of the extraction column will be loaded with different bio-crude compounds. As the bio-crude compounds are absorbed in scCO2 medium, solute-solute interaction effects will be negligible in this phase as compared to the liquid bio-crude phase. Exclusion of solute-solute interactions will simplify the system such that only solute-solvent binary interaction effects will now play the determining role in the phase behaviour description of supercritical extract phase. The application of stage-wise pressure reduction techniques for the purification of bio-compounds have been reported in the literature2, 14 but for mixtures other than bio-oils. A Peng-Robinson equation of state15 (PR-EOS) model was developed to investigate the phase behaviour of the solutes-rich supercritical phase. This model was subsequently validated against pilot plant scale trials. Another aim of this study is to first time compare the techno-economics of scCO2 separation of bio-crude with that of a conventional distillation process. This has been achieved by simulating both the SFE and distillation separation processes in Aspen Plus® and then evaluating the respective process economics.

2. Experimental methodology 2.1.

Materials

Carbon dioxide was purchased from Supagas (Australia), with purity ≥ 99.9 wt%. Biocrude was produced in-house from the HTL of black liquor, where the black liquor was a lignin-rich by-product of a bagasse pulping process. Phenol, p-cresol, catechol, 4ethylphenol, acetic acid, docosane, sulphuric acid and acetone were purchased from Sigma-Aldrich (Australia), each with purity ≥ 99.0 wt% except for sulphuric acid and 4ethylphenol which were ≥ 98.0 wt% and ≥ 97.0 wt% pure respectively.


Page 2 of 28

Page 3 of 28 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


2.2.

Bio-crude preparation and its characteristics

About 50 litres of bio-crude was produced from black liquor using the HTL continuous reactor facility at QUT. HTL liquefaction of the black liquor was performed at a temperature and pressure of 290oC and 75.8 bar respectively. The HTL reactor residence time was 60 minutes. The bio-crude product was stored at 2oC in a closed container prior to the SFE pilot plant extraction trials. The oil was homogenous, had a dense blackish appearance and a viscosity similar to water. The native HTL bio-crude had a pH of 9.0 but preliminary SFE pilot plant trials revealed that the extraction at such a high pH was problematic as it caused foaming, clogging and carry-over of water from the extraction column. To lower the pH of the bio-crude, sulphuric acid was incrementally added and then vigorously agitated with an electric mixer until a final pH of 4.4 was achieved. This pH-lowered bio-crude (pH=4.4) was centrifuged at 3300 rpm (Beckman GS-6R centrifuge, Marshall Scientific, USA) for 5 minutes, to remove precipitates and suspended solids.

2.3.

The SFE pilot plant setup

The SFE pilot plant used to determine the initial extraction and verify the predicted stagewise fractionation of bio-crude components is shown schematically in Figure 1. The pilot plant was purchased from Applied Separations (USA) and installed and commissioned at the QUT Pilot Plant Precinct. It consisted of a CO2 reservoir, bio-crude feed tank, CO2 preheater, temperature-regulated extraction column and two separators in series. Pressure in the separators was controlled with micrometering valves.

Figure 1. Pilot plant setup used in this work for supercritical extraction and fractionation of bio-crude (T: temperature control, Sep: separator, MV: micrometering valve). Sep-1 and Sep-2 were wrapped in trace heaters to compensate for the cooling effects resulting from depressurisation of the extract streams.



2.4.

Page 4 of 28

Extraction and Fractionation Procedure

Carbon dioxide from the reservoir cylinder is supplied at a set flow rate by a high pressure pneumatic pump (Haskel, USA). This high-pressure CO2 is then passed through a 1250 watts pre-heater to bring the CO2 up to the desired extraction temperature, before entering into the extraction column. The extraction column is a 4 litre stainless steel tubular vessel in which CO2 enters from bottom and bio-crude from top. The CO2 and biocrude streams flow in counter current over a densely packed bed made up of small tubular elements. The CO2 stream absorbs the majority of non-aqueous bio-crude and then leaves from the top of the extraction column where it is fed into the two separators in series. The remaining bio-crude and the majority of water is continuously drained from the bottom of the extraction column as raffinate. Shut-off and micrometering valves are positioned so as to produce the required pressures in column and separators respectively. Once the steady state operation has been reached and no more fluctuations in temperatures, pressures and flowrates are observed, sampling procedures are initiated. Separator fractions and column raffinate samples were collected every 15-30 minutes once continuous operation was achieved. From the bio-oil solubility in scCO2 reported in the literature3, 16 and preliminary trials on a lab scale solubility cell, it was determined that minimum mass flow ratios of CO2 to biooil of under 10 could be used, in the pilot plant trials, to ensure getting saturated extractions and consistent solubility data for analysis. Normally S/B should be greater than 10 to maximise yields but was limited in the pilot plant trials to less than this value because of pump cavitation issues. Run conditions used for the pilot plant extraction and fractionation trials are summarized in Table 1. Table 1. Parameters used in this work for the supercritical CO2 pilot plant extraction and fractionation of bio-crude produced from HTL of sugarcane bagasse black liquor. Extraction was performed at 55oC temperature and 206.4 bar pressure, and Sep-2 was maintained at 18.4oC temperature and 46.8 bar pressure.

Sep-1 No. 1 2 3 4 5 6 7 8 9 10 1

Press. (bar)

Temp. (oC)

137.6

49

116.3

47

91.5

43

CO2 flow1 (mL/min)

Bio-crude flow (mL/min)

S/B ratio2 (mass basis)

217 202 203 307 284 299 260 291 303 300

89 88 89 68 50 52 41 41 42 42

2.5 2.3 2.3 4.5 5.7 5.7 6.3 7.1 7.2 7.1

Flow rate is given for CO2 at extraction column inlet. Corresponding CO2 inlet temperature and pressure conditions were 48.6oC and 206.4 bar respectively. 2 Bio-crude density was 1.09 g/mL.

After reviewing the temperature and pressure conditions commonly found in the literature4, 16 for such an extraction process and to ensure the density difference between


Page 5 of 28 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


the two phases inside the extraction column was at least 150 g/L17 to avoid flooding, the conditions in Table 1 were chosen in this work to make a comparison between our experimental fractionation results and the model predictions. Maximum CO2 density used in the pilot plant trials was 763 g/L.

2.5.

Gas chromatography mass spectrometry (GC-MS) analysis

The quantities of several key compounds present in the bio-crude and extraction products were determined by GC-MS analysis. This process was performed on an Agilent (US) 6890 Series Gas Chromatograph and a HP 5975 mass spectrometer detector, employing helium as the carrier gas. The installed column was a dimethyl polysiloxane Agilent DB 5-MS, 30 m x 0.32 mm x 0.25 μm. A split-less injection of 2 μL was delivered to the injection port set at 250 °C. The temperature program commenced at 70 °C and was heated at a rate of 5 °C.min-1 to a temperature of 320 °C. Compounds were identified from the spectra by means of the Wiley library-HP G1035A and NIST mass spectra libraries and subsets-HP G1033A (a criteria quality value >90% was used). Analytical samples were prepared in acetone at a concentration of 0.05 mg/mL. Standard solutions of pure chemicals were also prepared in acetone, in order to produce a 5-point calibration curve over a concentration range of 0.025 to 0.3 mg/mL. All standards and analytical samples were spiked with Docosane at a concentration of 0.06 mg/mL, to act as an internal standard.

2.6.

Nuclear magnetic resonance (NMR) spectroscopy

Each sample (100 mg) of the collected oil fraction was dissolved in 0.9 mL of deuterated water (D2O)) and filtered. The 1H spectra were then recorded at 25 °C on a Bruker AVANCE III HD 600 MHz NMR spectrometer (Agilent, US) equipped with a cooled 5 mm TCI Cryoprobe. A total of 8 transients having an acquisition time of 1.7 seconds and a spectral width of 9 kHz were recorded using the Bruker pulse sequence noesygppr1d which features water suppression. The triplet phenol reference peak was used as an internal chemical shift reference point (δH = 7.25). Processing used shifted squared sine bell Gaussian apodization in 1H. Data processing and plots were carried out using ACD/NMR processing software, with automatic phase and baseline correction.

3. Thermodynamic modelling Modelling was implemented in Aspen Plus® software, using the Peng-Robinson-BostonMathias (PR-BM) property method.18 The Peng-Robinson Equation of State (PR-EOS)15 forms the basis of the PR-BM property method, and BM alpha function and asymmetric mixing rules are used in conjunction with the EOS to make it suitable for modelling polar, non-ideal chemical systems. Eqs 1-14 are mathematical expression of PR-BM model with asymmetric mixing rules. 𝑅𝑇

𝑎

𝑃 = 𝑉𝑚 ― 𝑏 ― 𝑉𝑚(𝑉𝑚 + 𝑏) + 𝑏(𝑉𝑚 ― 𝑏)

(1)

𝑏 = ∑𝑖𝑥𝑖𝑏𝑖

(2)



Page 6 of 28

𝑎 = 𝑎0 + 𝑎1

(3)

𝑎0 = ∑𝑖∑𝑗𝑥𝑖𝑥𝑗(𝑎𝑖𝑎𝑗)0.5(1 ― 𝑘𝑖𝑗)

(4)

Eq 4 is the standard quadratic mixing term, where 𝑘𝑖𝑗 has been made temperaturedependent (3) (2) 𝑘𝑖𝑗 = 𝑘(1) 𝑖𝑗 + 𝑘𝑖𝑗 𝑇 + 𝑘𝑖𝑗 𝑇

Where

(5) and superscripts (1), (2) and (3) are numbered terms in eq 5

𝑘𝑖𝑗 = 𝑘𝑗𝑖

[

𝑛

13 3

𝑛

1 𝑎1 = ∑𝑖 = 1𝑥𝑖 ∑𝑗 = 1𝑥𝑗((𝑎𝑖𝑎𝑗) 2𝑙𝑖,𝑗)

]

(6)

Eq 6 is an additional asymmetric term used to model highly non-linear systems (3) (2) 𝑙𝑖𝑗 = 𝑙(1) 𝑖𝑗 + 𝑙𝑖𝑗 𝑇 + 𝑙𝑖𝑗 𝑇

Where

(7)

𝑙𝑖𝑗 ≠ 𝑙𝑗𝑖 and superscripts (1), (2) and (3) are numbered terms in eq 7

The pure component parameters for PR-EOS are calculated as follows: 𝑅2𝑇2𝑐𝑖

𝑎𝑖 = 𝛼𝑖0.45724

(8)

𝑃𝑐𝑖

𝑅𝑇𝑐𝑖

(9)

𝑏𝑖 = 0.07780 𝑃𝑐𝑖

The parameter 𝛼𝑖 in Eq. 8 is used to improve the accuracy of predicted temperature response of the pure component vapour pressure. In standard PR-EOS, this parameter is expressed with eqs 10-11.

[

1

𝛼𝑖(𝑇) = 1 + 𝑚𝑖(1 ― 𝑇𝑟𝑖 2)

]

2

(10) (11)

𝑚𝑖 = 0.37464 + 1.54226𝜔𝑖 ―0.26992𝜔2𝑖

𝛼𝑖 defined in eq 10 is used when 𝑇𝑟 < 1 (subcritical temperature), otherwise Aspen BM alpha function (eqs 12-14) is used. 𝛼𝑖(𝑇) = [𝑒𝑥𝑝[𝐶𝑖(1 ― 𝑇𝑑𝑟𝑖)]]

2

(12)

𝑑𝑖 = 1 + 𝑚𝑖 2

(13)

𝐶𝑖 = 1 ― 1 𝑑𝑖

(14)

Binary interaction parameters (𝑘𝑖𝑗, 𝑙𝑖𝑗) must be determined from regression of phase equilibrium data. The optimized values of these binary interaction parameters were obtained by maximum-likelihood algorithm (eq 15), defined within the Aspen Plus® data regression system. 𝑁𝐷𝐺

𝑁𝑃

[(

𝑄 = ∑𝑛 = 1𝑤𝑛∑𝑖 = 1

𝑇𝑒,𝑖 ― 𝑇𝑚,𝑖 2 𝜎𝑇, 𝑖

) ( +

𝑃𝑒,𝑖 ― 𝑃𝑚,𝑖 2 𝜎𝑃, 𝑖

)

2 𝑁𝐶 ― 1 𝑥𝑒,𝑖,𝑗 ― 𝑥𝑚,𝑖,𝑗 𝜎𝑥, 𝑖,𝑗

+ ∑𝑗 = 1

(

)

2 𝑁𝐶 ― 1 𝑦𝑒,𝑖,𝑗 ― 𝑦𝑚,𝑖,𝑗 𝜎𝑦, 𝑖,𝑗

+ ∑𝑗 = 1

(

(15)


)]

Page 7 of 28 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


Table 2 provides the standard pure component properties of critical temperature (Tc), critical pressure (Pc) and acentric factor (ω), used in the Aspen Plus® modelling of the binary systems. Table 2. Critical properties of pure compounds used in the Aspen Plus® modelling of the binary systems

Component Carbon dioxide p-Cresol 4-Ethylphenol Phenol Catechol Acetic acid Water

Tc (oC)

Pc (bar)

ω

31.06

73.83

0.2236

431.5 443.3 421.1 490.85 318.8 373.9

51.5 42.9 61.3 74.9 57.9 220.6

0.5072 0.5154 0.4435 0.6937 0.4665 0.3449

The default binary interaction parameters available in in Aspen Plus® were adjusted in this study such that the PR-BM property method used in the analysis produced predictions which agreed more closely with experimental solubility data published in the open literature. Table 3 shows the deviations between the default Aspen Plus® predictions and experimental vapour-liquid equilibrium (VLE) data from literature, for all our binary systems. Regressed values of binary interaction parameters for all our binary systems are given in Table 4. Acetic acid in Aspen Plus® showed relatively poor agreement with experimental vapour phase solubility data giving an average absolute relative deviation (AARD) of about 30% when compared to Bamberger et al. (2000)19 and about 35% to Jonasson et al. (1998)20 data. On the other hand, liquid phase composition data of this system was reasonably represented with the same model, where the AARD between model predictions and both experimental studies19, 20 was within 10%. Bamberger et al. (2000)19 also pointed out towards difficulty in modelling the VLE data of acetic acid, whence his selected model represented the vapour phase composition with yet 18% deviation to experimental data, but only when more sophisticated modelling approach of taking into account the dimerization of acetic acid was adopted. Yet, the model predictions of Bamberger et al. (2000)19 were 50% smaller than reported by Jonasson et al. (1998)20. This means the model chosen in this work, and which represents all our other binary systems very well, can be reasonably extended to acetic acid and CO2 binary system too, as the average deviation between our model predictions and experimental data of different sources19, 20 is on average 25-35% AARD. For catechol experimental VLE data was not available, as catechol will be present in solid phase and will exhibit solid-fluid equilibrium at our interested supercritical extraction conditions. For this binary system no regression was done, and it was found that the default model predictions were in reasonable agreement to experimental solid-fluid data of Garcia et al. (2001)21, with average deviation of less than 20% AARD for data determined under 200 bar pressure. Similarly, no experimental VLE data was available for 4-ethylphenol and CO2 binary system, so no regression could be performed on this system as well, rendering the model description of this system totally predictive in nature based upon critical properties of pure components listed in Table 2 above.



Page 8 of 28

Table 3. Percent AARD between predicted and experimental VLE data for different solute-CO2 binary systems using the default regression coefficients for the PR-BM property method model available in Aspen Plus®

Binary system Phenol Catechol Acetic acid p-Cresol Water

Experimental data Pfohl et al. (1997)22 Yau et al. (1992)23 Garcia et al. (2001)21 Bamberger et al. (2000)19 Jonasson et al. (1998)20 Lee et al. (1999)24 Pfohl et al. (1997)22 Bamberger et al. (2000)19 Dohrn et al. (1993)25 Briones et al. (1987)26

Isotherms (Temp. in K) 373.15 348, 373, 398 333.15, 348.15, 363.15 353.2, 313.2, 333.2 323, 348 353.15, 393.15, 423.15 373.15 323.2, 333.2, 353.1 323.1 323.14

Model deviation (% AARD) 4.8 15.5, 8.4, 4.9 15.5, 21.2, 21.8 28.0, 34.8, 31.8 45.5, 24.5 4.2, 2.4, 1.9 9.1 0.9, 0.6, 0.4 1.0 1.5

Table 4. Numerical values of binary interaction parameters obtained after regressing the experimental VLE data (Table 3) of different solute-CO2 binary systems, with the EOS model of PR-BM property method within Aspen Plus® data regression system

Binary system Phenol Acetic acid p-Cresol Water

1 𝑘(1) 𝑖𝑗 0.08882 0.05469 0.29673 -0.32147 1

𝑘(2) 𝑖𝑗 -0.00057 0.001

𝑙(1) 𝑖𝑗 0.11836 0.18117 0.32347 -0.32052

𝑙(1) 𝑗𝑖 0.02185 0.06455 0.21729 0.19947

𝑙(2) 𝑖𝑗 -0.00065 -

𝑙(2) 𝑗𝑖 -0.00065 -

component i is solute and component j is CO2

4. Process design and techno-economic evaluation using Aspen Plus® After modelling the individual binary phase behaviour of each selected chemical compound with scCO2, simulations were run in Aspen Plus® to determine the potential for fractionation of the column extract stream. The solute-solute interaction parameters were set to zero. Only solute-solvent binary interaction parameters were employed to determine if the binary interaction parameters alone were sufficient to describe the phase behaviour and predict the fractionation characteristics for a defined column extract stream composition. In total, four process scenarios were simulated and economically evaluated. These scenarios are summarised in Table 5.


Page 9 of 28 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


Table 5. Description of Aspen Plus® simulation scenarios simulated in this work, for recovery of compounds from bio-crude.

Scenario  Process-1 (P-1)

  

Process-2 (P-2) Process-3 (P-3) Process-4 (P-4)

   

Initial scCO2 extraction of bio-crude from aqueous component; two-stage scCO2 fractionation of biocrude extract; further purification of fractionated components using conventional distillation and; catechol recovery from aqueous extraction column raffinate using multi-stage evaporation. As for P-1 but with single-stage scCO2 fractionation to recover column extract Atmospheric distillation of bio-crude; distillation includes the recovery and separation of the catechol and water components. As for P-3 but with multi-stage evaporation to recover catechol from the bottom stream of the first distillation column

A FLASH2 separator unit operation was employed in the Aspen Plus® simulation model to produce the solutes-rich stream representative of supercritical extract stream leaving our pilot plant extraction column. The governing relationships used in the FLASH2 model are not reported here as they are readily available in the open literature.27-29 Upon depressurization of the extract stream the predicted equilibrium composition of both the liquid and vapour phases were dictated by the thermodynamic models described in Section 3. Two additional FLASH2 units downstream of the solute rich extraction stream were used in the Aspen Plus® flowsheet to simulate the pilot plant separators. In the simulation the extraction column and separators were operated at the same pressure and temperature conditions maintained in the pilot plant trials. In the simulation it was assumed that downstream distillation of the extracted products would be used to recover individual bio-crude compounds. The RadFrac® unit available in Aspen Plus® was used to simulate the additional distillation columns. For comparison purposes Aspen Plus® simulation was also developed in which all bio-crude products were recovered through conventional distillation. By way of example Figure 2 is the process flowsheet of the SFE and fractionation sections of the P-1 scenario. Process flow sheets for the remaining scenarios are provided as Supporting Information (Figures S1-S4). Referring to Figure 2, bio-crude is pumped from ambient conditions (22oC, 1 bar) to 206.4 bar pressure while the CO2 is recycled from downstream units at 206.4 bar pressure and then preheated along with bio-crude in a preheater to 55oC.



Both are then flashed separated in an extraction column at 55oC temperature and 206.4 bar pressure. Temperature and pressure conditions for fractionation are selected by the model for maximum separation between catechol and the remaining compounds.

Figure 2. Aspen Plus® process flowsheet for supercritical extraction of bio-crude followed by two-stage fractionation of column extract (part of P-1).

4.1.

First separator

Figure 3 shows the effect of temperature on distribution coefficients (K) of different components in a typical SFE fractionation process. The distribution coefficient is the mole fraction of a component in supercritical CO2 phase divided by the mole fraction of that component in liquid phase. It is evident from Figure 3 that the effect of temperature on the extent of separation between mixture components is significant for all components except catechol and water. Of the selected compounds used in the current study therefore only catechol and water will be retained in the first separator upon depressurization. Figure 3 also indicates that the distribution coefficients for phenol, p-cresol, 4ethylphenol and acetic acid rise rapidly as the temperature drops below 45oC. In our first separator the cooling effect of depressurization was compensated for to some extent by external heat provision resulting in the temperature dropping to 43.1 oC. Without any external heat supply in the first separator, the temperature will drop to 39 oC and eventually the mixture will revert to a liquid phase with no feed going into second separator. At 43.1 oC, all components, except catechol and water, will have K values


Page 10 of 28

Page 11 of 28 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


greater than 5, corresponding to favourable separation process design conditions. Lowering the temperature further to 40oC will further increase the catechol K value by almost two-fold (1.8 times).

Figure 3. Effect of temperature on distribution coefficients of components to be fractionated by stagewise pressure reduction.

It is a well-reported phenomenon for supercritical extraction technologies that as the pressure increases, separation among components start to diminish,1, 30 as can be observed in Figure 4 where the separation factor (α) is the ratio of the distribution coefficient of one solute component relative to another.

Figure 4. Separation factors of components tend to decrease and approach unity at higher pressures.



By contrast, the lower the fractionation pressure, the greater the possible separation among components. However, there is a trade-off; inspection of the distribution coefficients in Figure 5 shows that as pressure decreases, so too do the distribution coefficients of components. For example, at 75 bar the α values are very attractive (Figure 4), but this is achieved at the expense of extremely low distribution coefficients (Figure 5).

Figure 5. Distribution coefficients of components will decrease with decrease in pressure.

A suitable compromise pressure condition can be found at 90 bar, where α values of all our product compounds relative to catechol, are acceptably high as are their respective distribution coefficients. The next least soluble compound after catechol and water is phenol, and at 90 bar its K value of 5.2 will drop to just 0.55 at 75 bar. For practical separation, an α value of at least 2 is necessary.30-32 In our extract mixture, though all compounds show quite higher α values in reference to catechol, even at pressures as high as 180 bar, but at such pressure the catechol K value is about 11 times greater than at 90 bar.

4.2.

Second separator

The second separator was operated at 60 bar and 32oC, to allow pressurized recycling of lean CO2 coming off it. It is important to keep the temperature in the second separator slightly above the saturation temperature of CO2 (22 oC) at 60 bar to keep most of CO2 in vapour form to be recycled. So basically what has been done in this work is provision of external heat supply in both separators so as to ultimately keep the second separator temperature at 32 oC. Through iteration on second separator, it was found that raising the temperature of the second separator from 22 oC to 32 oC will increase the pressurized recycling of CO2 from 11.8% to 93.9% of total CO2 in use.


Page 12 of 28

Page 13 of 28 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


4.3.

Recycling

Vapour CO2 leaving the second separator is cooled down to convert it to liquid form and then pumped to 206.4 bar again to introduce it at the preheater inlet for reuse in the extraction column. Liquid product fractions are collected from separators 1 and 2 (SEP1, SEP-2) in collectors 2 and 3 (COL-2, COL-3) respectively. Extraction column raffinate is collected in collector 1 (COL-1). A small amount of CO2, about 6% of total in use, is released from the liquid products recovered at ambient pressure from the column and both separators. It was calculated in this work that compression and recycling of this residual CO2 is more economical (Figure 6) than to make it up from an external supply of CO2. In Scenario P-2, single stage collection is performed at 60 bar and 32oC. For this scenario the amount of residual CO2 being recycled at ambient pressure is 3% of total CO2 in use. In the P-1 and P-2 scenarios, the amount of CO2 being recompressed from ambient conditions are 196 kmol/hr and 102.5 kmol/hr respectively.

Figure 6. Operating cost of CO2 compression from ambient to 60 bar (liquid state) pressure vs liquid CO2 make-up cost.

4.4.

Product purification

Liquid products from the double and single separators in P-1 and P-2 respectively are sent to distillation columns where further separation and purification takes place. For both of these scenarios the extraction column raffinate is fed to a multi-stage evaporator set for single product (catechol) recovery (Figure 7). In P-1 and P-2, raffinate from the extraction column contains predominantly unrecovered catechol (86 wt% of catechol initially in the bio-crude feed) and water. This catechol/ water mixture is assumed in these scenarios to be passed through a multi-stage evaporation process to recover the catechol. A four stage pressure reduction and vapour heat recovery regime was implemented in the multi-stage evaporator set with the final (predominantly water) component being condensed at 0.065 bar (abs) and 40.3oC. The multi-stage evaporator set was predicted to remove 88 wt% of water in the extraction



column raffinate. A final distillation step is performed post-evaporation to remove the remaining water and recover commercial purity catechol.


Page 14 of 28

Page 15 of 28 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 7. Aspen Plus® flowsheet for the multi-stage evaporation and distillation processes used in the recovery of products following scCO2 extraction and fractionation (Scenario P-1)



Page 16 of 28

Modelling conventional distillation of the biocrude component of the scCO2 extraction products in P-1 and P-2 predicted separation and recovery of (in order of ascending boiling point temperature) water, acetic acid, phenol, p-cresol, 4-ethylphenol and catechol. The bottoms stream of the first distillation columns in P-3 and P-4 each contain a dilute solution of catechol in water. Simulating the recovery of catechol from the bottoms stream by subsequent conventional distillation (P-3) and multi-stage evaporation (P-4) enables a direct techno-economic comparison to be made between these two options for catechol recovery. Steam was assumed as the heating medium in the extraction, fractionation, distillation and evaporation stages of the proposed process scenarios. Wherever the distillation top temperature was more than 140oC, heat recovery was used for steam generation. Heat recovered from streams less than 140oC was used for pre-heating the bio-crude feed stream prior to distillation. Where appropriate distillate fractions were sent for further cooling and crystallisation to get the final products in market-ready form.

4.5.

A techno-economic assessment of process scenarios

The proprietary Aspen Process Economic Analyzer® was used with Aspen Plus® process simulation software to undertake a techno-economics assessment of the four process design scenarios. The compositional analysis of the bio-crude used in the simulation was based on an analysis of bio-crude produced during HTL continuous reactor pilot plant trials at QUT. The HTL feedstock used in these trials was a lignin-rich black liquor produced from the bio-refining of bagasse. The identity and relative concentrations of five main chemicals in the bio-crude were determined using GC-MS and NMR analysis. A normalised relative bio-crude composition based on these five chemicals (Table 6) were used as inputs in the simulation model. An S/B mass ratio of 6.2 was used in our simulation work. The effects of higher S/B ratios (of up to 20) are discussed in the technoeconomic assessment of the P-1 scenario. Table 6. Composition of bio-crude used in Aspen Plus® simulations of this work

Component

Phenol

p-Cresol

4-Ethyl phenol

Catechol

Acetic Acid

Water

Composition (wt%, normalized)

0.88

0.03

0.03

0.44

8.61

90

Bio-crude was the primary raw material input to all the simulated process scenarios and its value was assumed to be defined by its heating value (3717 kJ/kg) relative to crude oil and current crude oil prices.33 Unit bio-crude production costs and end-product sales prices in this work are listed in Table 7.


Page 17 of 28 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


Table 7. Raw material cost and product prices used in techno-economic evaluations of this work

Material Bio-crude33, 34 Phenol36 Catechol38 p-Cresol40 1

Price (USD/tonne) 42 97100 353770 37630

Material 4-Ethylphenol35 Acetic Acid37 Aqueous1 (AA)39 -

Price (USD/tonne) 79560 23600 110 -

Here aqueous (AA) is a 30 mol% acetic acid solution in water

All product prices are based on 99% purity, except for 4-ethylphenol and aqueous (AA) which are 98% and 30% pure respectively. Economics were evaluated assuming a total plant life of 20 years and for a company hurdle rate of 20%. The plant start-up time was given as 18 months and the plant availability to be 95% (8327 hr). Results are presented in terms of costs associated with total capital, raw materials (feedstock), utilities and operating, as well as annual product sales and profit.

5. Results and Discussion The bio-crude pH was lowered from 9.0 to 4.4 by the addition of 2% (vol/vol) of 98 wt% pure sulphuric acid resulting in approximately 0.9 wt% of initial bio-crude dropping out of solution. The bio-crude was then centrifuged to remove any suspended solids (Figure 8) prior to extraction and fractionation in the scCO2 extraction pilot plant.

Figure 8. Black liquor bio-crude before (A) and after (B) acidification.

Sampling of extraction and fractionation pilot plant products were performed under steady state operating conditions. Throughout the pilot plant trials maximum standard deviations in temperature and pressure of the extraction column, separator-1 and separator-2 were ±0.5 oC, ±0.8 oC and ±1.1 oC, and ±2.9 bar, ±1.7 bar and ±1.4 bar respectively. Average extract yield was 1.0 wt% of bio-crude feed rate, and it varied over 0.4 wt% to 1.7 wt% for an S/B range of 2.3 to 7.2. The average mass fraction solubility of bio-crude in scCO2 was 0.00213 with a maximum relative standard deviation of 23.3%. This level of deviation in bio-crude solubility in scCO2 was deemed to be a reasonable indicator that equilibrium conditions had been achieved within the extraction column.



The relative concentrations of phenol, p-cresol, catechol and 4-ethylphenol were quantified using GC-MS and acetic acid concentration was measured using NMR. The volume of sample collected from Separator-2 during trials was sufficient to obtain triplicate GC-MS results although only sufficient to produce duplicate NMR measurements for acetic acid. Both the experimental and simulated results indicated that only catechol would be recovered from Separator-1. The concentration of catechol (0.142 mg/mL) in our bio-crude was almost half of the most abundant compound phenol (0.288 mg/mL); the polar nature catechol would mitigate against its extraction with scCO2. The sample volumes collected at Separator-1 samples were small compared to those collected at Separator-2. Also, the concentration of catechol in Separator-1 samples was not much, so the GC-MS results for separator-1 samples should be regarded here more of a qualitative nature. Analysis of Separator-2 samples indicated that the relative standard deviation in concentration measurements for phenol, p-cresol and 4-ethylphenol ranged from 14.4% to 24.8%, while for catechol and acetic acid the maximum deviation was 12.1%. Figure 9 shows a comparison between the pilot plant experimental relative concentrations determined by a GC-MS method compared with model predictions for Separator-2 extracts. Figure 10a is a comparison of model and experimental results determined both by GC-MS (phenol) and NMR (acetic acid), also for Separator-2 samples. Figure 10b shows a comparison of the predicted and measured (GC-MS) relative concentrations of catechol and p-cresol for Separator-1 samples.

Figure 9. Relative concentrations of compounds in Separator-2 samples of supercritical extract, collected at a temperature of 18.4 oC and a pressure of 46.8 bar. Legend numerical values correspond to first separator pressure conditions (in bar abs). Concentration measurements were determined by GC-MS; Aspen Plus® model PR-BM was used in the simulations.


Page 18 of 28

Page 19 of 28 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 10. Comparison of experimental scCO2 fractionation of extracted bio-crude with Aspen Plus® model of this work. (A) Data of phenol (GC-MS) and acetic acid (NMR) for fraction-2. (B) Catechol relative concentration in fraction-1 relative to p-cresol in the same fraction. Legend numerical values in both figures (A) and (B) correspond to first separator pressure conditions. Fraction-2 was collected at 18.4 oC temperature and 46.8 bar pressure.

Phenol and acetic acid were relatively abundant compounds of bio-crude; other compounds were found to be present in comparatively small quantities. Figures 9 and 10 indicate reasonable qualitative agreement between experimental and model data for all compounds. However quantitative comparison in terms of relative concentrations was good for phenol and acetic acid only due to the relative abundance (and therefore reduced experimental uncertainty) associated with these two compounds. Absolute deviation between experimental and model data for phenol ranged from 17.6% to 20.4%, and for acetic acid it was 31.0%. By comparing the measured and modelled component mass ratios some of the experimental uncertainty associated with measuring the small quantities of p-cresol, 4ethylphenol and catechol present in the bio-crude fractions, can be circumvented. (Figure 11).



Figure 11. Mass ratios of compounds in second fraction of supercritical extract, collected at 18.4 oC temperature and 46.8 bar pressure. Legend numerical values correspond to first separator pressure conditions. Amounts determined by GC-MS method. Aspen Plus® model PR-BM was used in simulation.

Inspection of Figure 11 indicates reasonable agreement in terms of the mass ratios between experimental and model data for these more minor components. Maximum absolute deviation between model and experimental data for catechol/p-cresol and catechol/4-ethylphenol was under 20% (ranging from 1.21% to 19.61%) except for samples collected at 91.5 bar separator-1 pressure for which absolute deviation reached 44.81% and 48.63% respectively. This discrepancy was most probably caused by catechol precipitation during the fractionation process, as inspection of Figure 11 indicates that the ratio of 4-ethylphenol/p-cresol (where catechol and the potential for crystallisation is absent) showed a maximum absolute deviation of just 9.98% to the model, at all studied conditions. Figure 5 indicates that extent of fractionation of the bio-crude extract into two fractions, by means of stage-wise pressure reduction is limited by the respective phase equilibrium characteristics of the bio-crude components. Extraction can be more effective into distinct fractions in the first place in column where K values are more favourable, and should be above 1 for a practical separation.31, 32 Such higher K values, for some compounds, in an extraction column become possible due to involvement of solute-solute interactions and tendency of being selectively extracted into vapour phase in comparison to other compounds. For example, in Figure 5, acetic acid is showing higher K values than many others in a supercritical extract stream, and comes out potentially as a good candidate when to be fractionated into a lower pressure separator, but when it is seen in the context of supercritical extraction itself in a column it is well-known that acetic acid shows very small K value, about 0.03 (weight basis),2 and largely remains in liquid (water) phase. It means, though, a compound like acetic acid might be a bad choice when it comes to extracting it , but once extracted out of bio-crude, this compound shows better tendency to be further fractionated by stage-wise pressure reduction. Our initial pilot plant runs on a different kind of bio-crude also endorsed the possibility of such a stage-


Page 20 of 28

Page 21 of 28 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


wise fractionation in which acetic acid was being collected in the last separator, just like the simulation results of this work suggested. The only difference between simulations of P-1 and P-2 processes was that in former twostage fractionation was done based on optimal conditions of our model, while in later only single stage fractionation was performed. The effect of this fractionation will be seen translated into overall techno-economics of process, presented later in this document. Column extract yields in P-1 and P-2 were 12.95 wt% and 12.98 wt% respectively, whereby 9.15 wt% and 9.39 wt% respectively of extracted products were recycled back from separator-2 to the extraction column. Recycle stream from separator-2 contained primarily water along with small amounts of acetic acid, therefore it was deemed not economical in this work to remove these two components from recycled CO2 before putting it back into the extraction column. The final collected product yields of both these P-1 and P-2 processes were 9.59 wt% each (dry basis), with 2.03 wt% and 1.87 wt% respectively of feed bio-crude water contents in them. More than 94% of CO2 being used in both processes was recycled off sepearator-2 at 60 bar, the rest being recycled at ambient pressure after depressurization of liquid products from the extraction column and both separators. In all four process designs, three products including acetic acid, catechol and phenol were produced with at least 99% purity. On the other hand, 4-ethylphenol, p-cresol and weak acetic acidic solutions in water (aqueous AA) were produced in purity ranges of 80%85.5%, 78.3%-85% and 21%-26% respectively. Water/acetic acid and p-cresol/4ethylphenol could not be separated from each other beyond the above mentioned ranges. Multiple stage evaporation recovered 97.5% of catechol in the P-1, P-2 and P-4 scenarios and was able to remove 87.9% of the water entering the evaporator station. The performance of the multiple stage evaporation unit was evaluated in terms of tonnes of water evaporated per tonne of steam supplied. This ratio was 3.01 in both in P-1 and P2, and 4.58 in P-4. In P-4, the ratio was greater than 4 (ideal) because the feed stream into the evaporation station was at a higher temperature of 101oC than in P-1 and P-2 scenarios where it was 54.3oC. Final predicted recoveries of each product is shown in Figure 12 for all four process scenarios simulated. With the exception of acetic acid, the product recovery is similar for all four process scenarios. In the case of acetic acid P-1 shows improved recovery relative to the other process scenarios. This improved performance for the P-1 process scenario is attributed to 2-stage scCO2 fractionation where residual water removed at the first fractionation stage reduces the loss of acetic acid in the subsequent distillation stage (water and acetic acid have similar boiling points).



Figure 12. Compound recoveries of bio-crude into pure chemical products.

The results of the economic analyses carried out for all our four process scenarios are summarized in Figure 13. Some points of note are:      





Total Raw Material Costs are dominated by those associated with the purchase of the bio-crude feedstock (and are therefore identical for all scenarios); process P-3 is the most capital intensive scenario due to the high distillation capacity required; P-1 has a marginally higher Total Capital Cost than P-2 to accommodate a separate distillation unit to separate catechol from fraction-1; exceptionally high Total Utilities and Total Operating Costs are incurred in the P3 scenario due to the high steam required for distillation; P-4 has the lowest Total Utilities and Operating Costs due to the energy efficiency of water removal by the multiple stage evaporation unit; P-1 indicates the highest Total Product Sales due to the efficiency of product recovery of this scenario and the highest overall profitability generating annual Total Profits that are 7.49%, 10.93% and 10.60% greater than P-2, P-3 and P-4 respectively. Inclusion of costs incurred by the addition of sulphuric acid in bio-crude to lower pH in SFE process will not have a significant effect on the annual Total Profits of P-1 and P-2 scenarios as it will be an increase of just 23.8% in Raw Material Costs relative to those scenarios in which there had been no pH adjustment. When S/B ratios of 12.4 and 20 were used in the P-1 scenario, the Total Operating Cost increased by 8.77% and 16.71% respectively, while the annual Total Profits decreased by just 0.65% and 1.32% respectively relative to that produced for the base case condition of S/B = 6.2.


Page 22 of 28

Page 23 of 28 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 13. Techno-economic summary of four process simulations to compare basically supercritical separation of bio-crude with that of distillation.

6. Conclusions This study has used pilot plant data and process modelling to investigate the industrial scale use of scCO2 extraction and fractionation of bio-crude for the recovery of renewable chemicals from a lignin-rich HTL bio-crude. It was confirmed through pilot plant extraction and fractionation trials that it is possible to effectively model the extraction characteristics of a multi-component bio-crude by a series of individual solute-solvent binary interaction parameters regressed from experimental VLE data.



Aspen Plus® process and economic models of four design scenarios were developed to compare supercritical extraction with conventional distillation of bio-crude. These models indicated that two stage scCO2 extraction of bio-crude combined with multiple stage evaporation to remove water and recover catechol (process scenario P-1) generates 10.9% and 10.6% more Total Profit annually than distillation alone (P-3) and distillation with multiple stage evaporation (P-4) process scenarios respectively. It was established that that the improved economics of the scCO2 technology relative to other technology scenarios was maintained even working with S/B ratios as high as 20. These results suggest two main areas of future investigation to further improve the profitability of industrial scale scCO2 recovery of chemicals from bio-crude: 1. HTL production of bio-crude should be tailored to produce fewer compounds but in large amounts rather than more compounds in small amounts. This will simplify the post-extraction treatment for further purification of products; and 2. modelling of the scCO2 extraction (i.e. pre-fractionation) process itself is needed to identify extraction conditions with improved yields and product composition profiles.

Glossary and Nomenclature Model = Aspen Plus® PR-BM property method 𝑎𝑖, 𝑏𝑖 = model parameters for pure components 𝑎, 𝑏 = model parameters for mixture e = estimated data i = data for data point i, (eq 15) j = fraction data for component j (eq 15) 𝑘𝑖𝑗, 𝑙𝑖𝑗 = binary interaction parameters in model m = measured data NDG = the number of data groups in the regression case NC = the number of components present in the data group NP = the number of points in data group n P = pressure 𝑃𝑐 = critical pressure of a component Q = maximum-likelihood objective function to be minimized R = gas constant T = temperature 𝑇𝑐 = critical temperature of a component 𝑇𝑟 = reduced temperature Wn = the weight of data group n x, y = liquid and vapor mole fractions respectively 𝛼 = temperature function in eq 8


Page 24 of 28

Page 25 of 28 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


σ = standard deviation of the indicated data 𝜔 = acentric factor of a component

Associated Content Supporting Information Aspen Plus® process flowsheets for supercritical extraction and distillation processes (P2, P-3 and P-4), table listing summary of economic evaluation for different separation and purification processes of bio-crude (P-1 to P-4)

Author Information Corresponding Author *E-mail: [email protected]

Notes The authors declare no competing financial interest.

Acknowledgements This work was undertaken with Australian Federal Government and Queensland University of Technology support under the Australia-India Strategic Research Fund program.

References (1) Martinez, J. L. Supercritical fluid extraction of nutraceuticals and bioactive compounds; CRC Press: Boca Raton, FL, 2008. (2) McHugh, M. A.; Krukonis, V. J. Supercritical fluid extraction principles and practice; Butterworth-Heinemann: Boston, 1994. (3) Chan, Y. H.; Yusup, S.; Quitain, A. T.; Chai, Y. H.; Uemura, Y.; Loh, S. K. Extraction of palm kernel shell derived pyrolysis oil by supercritical carbon dioxide: Evaluation and modeling of phenol solubility. Biomass Bioenergy 2018, 116, 106-112. (4) Chan, Y. H.; Yusup, S.; Quitain, A. T.; Uemura, Y.; Loh, S. K. Fractionation of pyrolysis oil via supercritical carbon dioxide extraction: optimization study using response surface methodology (RSM). Biomass Bioenergy 2017, 107, 155-163. (5) Cheng, T.; Han, Y.; Zhang, Y.; Xu, C. Molecular composition of oxygenated compounds in fast pyrolysis bio-oil and its supercritical fluid extracts. Fuel 2016, 172, 49-57. (6) Feng, Y.; Meier, D. Extraction of value-added chemicals from pyrolysis liquids with supercritical carbon dioxide. J. Anal. Appl. Pyrolysis 2015, 113, 174-185. (7) Feng, Y.; Meier, D. Comparison of supercritical CO2, liquid CO2, and solvent extraction of chemicals from a commercial slow pyrolysis liquid of beech wood. Biomass Bioenergy 2016, 85, 346-354. (8) Mudraboyina, B. P.; Fu, D.; Jessop, P. G. Supercritical fluid rectification of lignin microwavepyrolysis oil. Green Chem. 2015, 17, 169-172. (9) Naik, S.; Goud, V. V.; Rout, P. K.; Dalai, A. K. Supercritical CO2 fractionation of bio-oil produced from wheat-hemlock biomass. Bioresour. Technol. 2010, 101, 7605-7613.



(10) Patel, R. N.; Bandyopadhyay, S.; Ganesh, A. Extraction of cardanol and phenol from bio-oils obtained through vacuum pyrolysis of biomass using supercritical fluid extraction. Energy 2011, 36, 1535-1542. (11) Perez, E.; Tuck, C. O.; Poliakoff, M. Valorisation of lignin by depolymerisation and fractionation using supercritical fluids and conventional solvents. J. Supercrit. Fluids 2018, 133, 690-695. (12) Rout, P. K.; Naik, M. K.; Dalai, A. K.; Naik, S. N.; Goud, V. V.; Das, L. M. Supercritical CO2 fractionation of bio-oil produced from mixed biomass of wheat and wood sawdust. Energy Fuels 2009, 23, 6181-6188. (13) Wang, J.; Cui, H.; Wei, S.; Zhuo, S.; Wang, L.; Li, Z.; Yi, W. Separation of Biomass Pyrolysis Oil by Supercritical CO2 Extraction. Smart Grid Renewable Energy 2010, 01, 98-107. (14) Arai, Y.; Sako, T.; Takebayashi, Y. Supercritical fluids: molecular interactions, physical properties and new applications; Springer: Berlin, 2013. (15) Peng, D.-Y.; Robinson, D. B. A new two-constant equation of state. Ind. Eng. Chem. Fundam. 1976, 15, 59-64. (16) Maqbool, W.; Hobson, P.; Dunn, K.; Doherty, W. Supercritical carbon dioxide separation of carboxylic acids and phenolics from bio-oil of lignocellulosic origin: Understanding bio-oil compositions, compound solubilities, and their fractionation. Ind. Eng. Chem. Res. 2017, 56, 31293144. (17) Machado, N. T. Fractionation of PFAD-compounds in countercurrent columns using supercritical carbon dioxide as solvent; Hamburg University of Technology: Germany, 1998. (18) PR-BM property method, Aspen Plus V8.4 Help, AspenTech. (19) Bamberger, A.; Sieder, G.; Maurer, G. High-pressure (vapor+ liquid) equilibrium in binary mixtures of (carbon dioxide+ water or acetic acid) at temperatures from 313 to 353 K. J. Supercrit. Fluids 2000, 17, 97-110. (20) Jónasson, A.; Persson, O.; Rasmussen, P.; Soave, G. S. Vapor–liquid equilibria of systems containing acetic acid and gaseous components. Measurements and calculations by a cubic equation of state. Fluid Phase Equilib. 1998, 152, 67-94. (21) García-González, J.; Molina, M. J.; Rodríguez, F.; Mirada, F. Solubilities of phenol and pyrocatechol in supercritical carbon dioxide. J. Chem. Eng. Data 2001, 46, 918-921. (22) Pfohl, O.; Avramova, P.; Brunner, G. Two-and three-phase equilibria in systems containing benzene derivatives, carbon dioxide, and water at 373.15 K and 10–30 MPa. Fluid Phase Equilib. 1997, 141, 179-206. (23) Yau, J. S.; Tsai, F. N. Solubility of carbon dioxide in phenol and in catechol. J. Chem. Eng. Data 1992, 37, 141-143. (24) Lee, M.-J.; Kou, C.-F.; Cheng, J.-W.; Lin, H.-m. Vapor–liquid equilibria for binary mixtures of carbon dioxide with 1, 2-dimethoxybenzene, 2-methoxyphenol, or p-cresol at elevated pressures. Fluid Phase Equilib. 1999, 162, 211-224. (25) Dohrn, R.; Bünz, A.; Devlieghere, F.; Thelen, D. Experimental measurements of phase equilibria for ternary and quaternary systems of glucose, water, CO2 and ethanol with a novel apparatus. Fluid Phase Equilib. 1993, 83, 149-158. (26) Briones, J.; Mullins, J.; Thies, M.; Kim, B.-U. Ternary phase equilibria for acetic acid-water mixtures with supercritical carbon dioxide. Fluid Phase Equilib. 1987, 36, 235-246. (27) Ahmed, T. Working Guide to Vapor-Liquid Phase Equilibria Calculations; Elsevier: Amsterdam, 2009. (28) Schefflan, R. Teach yourself the basics of Aspen Plus; John Wiley & Sons: Hoboken, New Jersey, 2017. (29) Tassios, D. P. Applied chemical engineering thermodynamics; Springer-Verlag: Berlin Heidelberg GmbH, 1993. (30) Lim, C.; Manan, Z.; Sarmidi, M. Simulation modeling of the phase behavior of palm oil-supercritical carbon dioxide. J. Am. Oil Chem. Soc. 2003, 80, 1147-1156. (31) Hrnčič, M. K.; Cör, D.; Verboten, M. T.; Knez, Ž. Application of supercritical and subcritical fluids in food processing. Food Qual. Saf. 2018, 2, 59-67.


Page 26 of 28

Page 27 of 28 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


(32) Gutiérrez, C.; Rodríguez, J. F.; Gracia, I.; de Lucas, A.; García, M. T. Modeling the phase behavior of essential oils in supercritical CO2 for the design of a countercurrent separation column. Ind. Eng. Chem. Res. 2014, 53, 12830-12838. (33) https://www.eia.gov/todayinenergy/prices.php, accessed: 07 Sep 2018. (34) https://www.eia.gov/energyexplained/index.php?page=about_energy_units, accessed: 07 Sep 2018. (35) https://www.sigmaaldrich.com/catalog/substance/4ethylphenol1221612307911?lang=en&re gion=AU, accessed: 03 Aug 2018. (36) https://www.sigmaaldrich.com/catalog/substance/phenol941110895211?lang=en®ion=AU , accessed: 03 Aug 2018. (37) https://www.sigmaaldrich.com/catalog/substance/aceticacid60056419711?lang=en®ion= AU, accessed: 03 Aug 2018. (38) https://www.sigmaaldrich.com/catalog/substance/pyrocatechol1101112080911?lang=en® ion=AU, accessed: 03 Aug 2018. (39) https://www.indiamart.com/proddetail/dilute-acetic-acid-9837084655.html, accessed: 03 Aug 2018. (40) https://www.sigmaaldrich.com/catalog/substance/pcresol1081410644511?lang=en®ion= AU, accessed: 03 Aug 2018.

For Table of Contents Only




Page 28 of 28

Extraction and purification of renewable chemicals from hydrothermal

Recommend Documents