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Maqbool, Wahab, Dunn, Kameron, Doherty, William, McKenzie, Neil,Cronin, Dylan, & Hobson, Philip(2019)Extraction and purification of renewable chemicals from hydrothermal liq-uefaction bio-oil using supercritical carbon dioxide: A techno-economicevaluation.Industrial & Engineering Chemistry Research, 58(13), pp. 5202-5214.

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https://doi.org/10.1021/acs.iecr.8b05366

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: p.hobson@qut.edu.au

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, the internal rate of return (IRR) value was

similar for both SFE and distillation combined with multistage evaporation processes.

Solvent/bio-oil (S/B) ratio will have considerable impact on total profits of SFE process

in relation to distillation combined with multistage evaporation.

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

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. Bio-

oils 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%. Bio-

crude 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, 4-

ethylphenol, acetic acid, docosane, sulphuric acid and acetone were purchased from

Sigma-Aldrich (Australia), each with purity ≥ 99.0 wt% except for sulphuric acid and 4-

ethylphenol which were ≥ 98.0 wt% and ≥ 97.0 wt% pure respectively.

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 stage-

wise 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 pre-

heater, 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. 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 bio-

crude 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 bio-

oil 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.

No. Sep-1

CO2 flow1 (mL/min)

Bio-crude flow (mL/min)

S/B ratio2 (mass basis)

Press. (bar)

Temp. (oC)

1 137.6 49

217 89 2.5 2 202 88 2.3 3 203 89 2.3 4

116.3 47 307 68 4.5

5 284 50 5.7 6 299 52 5.7 7

91.5 43

260 41 6.3 8 291 41 7.1 9 303 42 7.2

10 300 42 7.1 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

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-Boston-

Mathias (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.

𝑃 =𝑅𝑇

𝑉𝑚−𝑏−

𝑎

𝑉𝑚(𝑉𝑚+𝑏)+𝑏(𝑉𝑚−𝑏) (1)

𝑏 = ∑ 𝑥𝑖𝑏𝑖𝑖 (2)

𝑎 = 𝑎0 + 𝑎1 (3)

𝑎0 = ∑ ∑ 𝑥𝑖𝑥𝑗(𝑎𝑖𝑎𝑗)0.5

(1 − 𝑘𝑖𝑗)𝑗𝑖 (4)

Eq 4 is the standard quadratic mixing term, where 𝑘𝑖𝑗 has been made temperature-

dependent

𝑘𝑖𝑗 = 𝑘𝑖𝑗(1)

+ 𝑘𝑖𝑗(2)

𝑇 + 𝑘𝑖𝑗(3)

𝑇⁄ (5)

Where 𝑘𝑖𝑗 = 𝑘𝑗𝑖 and superscripts (1), (2) and (3) are numbered terms in eq 5

𝑎1 = ∑ 𝑥𝑖[∑ 𝑥𝑗((𝑎𝑖𝑎𝑗)1 2⁄ 𝑙𝑖,𝑗)1 3⁄𝑛𝑗=1 ]

3𝑛𝑖=1 (6)

Eq 6 is an additional asymmetric term used to model highly non-linear systems

𝑙𝑖𝑗 = 𝑙𝑖𝑗(1)

+ 𝑙𝑖𝑗(2)

𝑇 + 𝑙𝑖𝑗(3)

𝑇⁄ (7)

Where 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:

𝑎𝑖 = 𝛼𝑖0.45724𝑅2𝑇𝑐𝑖

2

𝑃𝑐𝑖 (8)

𝑏𝑖 = 0.07780𝑅𝑇𝑐𝑖

𝑃𝑐𝑖 (9)

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.

𝛼𝑖(𝑇) = [1 + 𝑚𝑖(1 − 𝑇𝑟𝑖1 2⁄

)]2 (10)

𝑚𝑖 = 0.37464 + 1.54226𝜔𝑖 − 0.26992𝜔𝑖2 (11)

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 − 𝑇𝑟𝑖𝑑)]]

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.

𝑄 = ∑ 𝑤𝑛 ∑ [(𝑇𝑒,𝑖−𝑇𝑚,𝑖

𝜎𝑇,𝑖)

2

+ (𝑃𝑒,𝑖−𝑃𝑚,𝑖

𝜎𝑃,𝑖)

2

+ ∑ (𝑥𝑒,𝑖,𝑗−𝑥𝑚,𝑖,𝑗

𝜎𝑥,𝑖,𝑗)

2𝑁𝐶−1𝑗=1 + ∑ (

𝑦𝑒,𝑖,𝑗−𝑦𝑚,𝑖,𝑗

𝜎𝑦,𝑖,𝑗)

2𝑁𝐶−1𝑗=1 ]𝑁𝑃

𝑖=1𝑁𝐷𝐺𝑛=1

(15)

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 Tc (oC) Pc (bar) ω Carbon dioxide

31.06 73.83 0.2236

p-Cresol 431.5 51.5 0.5072 4-Ethylphenol 443.3 42.9 0.5154 Phenol 421.1 61.3 0.4435 Catechol 490.85 74.9 0.6937 Acetic acid 318.8 57.9 0.4665 Water 373.9 220.6 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.

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

Experimental data Isotherms (Temp. in K) Model

deviation (% AARD)

Phenol Pfohl et al. (1997)22 373.15 4.8 Yau et al. (1992)23 348, 373, 398 15.5, 8.4, 4.9

Catechol Garcia et al. (2001)21 333.15, 348.15, 363.15 15.5, 21.2, 21.8

Acetic acid Bamberger et al. (2000)19 353.2, 313.2, 333.2 28.0, 34.8, 31.8 Jonasson et al. (1998)20 323, 348 45.5, 24.5

p-Cresol Lee et al. (1999)24 353.15, 393.15, 423.15 4.2, 2.4, 1.9 Pfohl et al. (1997)22 373.15 9.1

Water Bamberger et al. (2000)19 323.2, 333.2, 353.1 0.9, 0.6, 0.4 Dohrn et al. (1993)25 323.1 1.0 Briones et al. (1987)26 323.14 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 𝑘𝑖𝑗(1)1 𝑘𝑖𝑗

(2) 𝑙𝑖𝑗

(1) 𝑙𝑗𝑖

(1) 𝑙𝑖𝑗

(2) 𝑙𝑗𝑖

(2)

Phenol 0.08882 - 0.11836 0.02185 - -

Acetic acid 0.05469 - 0.18117 0.06455 - -

p-Cresol 0.29673 -0.00057 0.32347 0.21729 -0.00065 -0.00065

Water -0.32147 0.001 -0.32052 0.19947 - - 1 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.

Table 5. Description of Aspen Plus® simulation scenarios simulated in this work, for recovery of

compounds from bio-crude.

Scenario

Process-1 (P-1)

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.

Process-2 (P-2) As for P-1 but with single-stage scCO2 fractionation to recover

column extract

Process-3 (P-3) Atmospheric distillation of bio-crude; distillation includes the recovery and separation of the

catechol and water components.

Process-4 (P-4) 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, 4-

ethylphenol 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

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 stage-

wise 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.

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 (SEP-

1, 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.

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)

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 techno-

economic 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, 34 Unit bio-crude production costs and end-product sales

prices 35 in this work are listed in Table 7.

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

Material Price (USD/tonne) Material Price (USD/tonne) Bio-crude 42 4-Ethylphenol 1200

Phenol 1500 Acetic Acid 1200 Catechol 2400 Aqueous1 (AA) 360 p-Cresol 1100 - -

1 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 10%. 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.

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, 4-

ethylphenol 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-

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 two-

stage 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/4-

ethylphenol 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 P-

2, 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 P-

3 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 16.7% and 11.8% greater than P-2 and P-4 respectively. P-3

did not produce any profit, rather gave a negative value for annual total profits of

about 6.3 million USD. At base plant capacity of 22.8 tonne/hr, however the IRR

values of 15.0%, 14.7%, -2.1% and 15.3% were obtained for P-1, 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 20.2 were used in the P-1 scenario, the total operating

cost increased by 9.1% and 17.4% respectively, while the annual total profits

decreased by 22.7% and 47.8% respectively relative to that produced for the base

case condition of S/B = 6.2. For P-1, when S/B ratio of 6.2 was increased to 12.4

and 20.2, the corresponding IRR were 12.3% and 9.5% respectively.

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

16.7% and 11.8% more total profit annually than SFE single stage (P-2) and distillation

combined with evaporation (P-4) respectively. Distillation alone (P-3) did not produce

any profit, rather gave a negative value for annual total profits of about 6.3 million USD.

At base plant capacity of 22.8 tonne/hr, however the IRR values of 15.0%, 14.7%, -2.1%

and 15.3% were obtained for P-1, P-2, P-3 and P-4 respectively. Solvent/biocrude ratio

will have considerable impact on total profits of SFE process in relation to distillation

combined with multistage evaporation.

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

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 (P-

2, 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: p.hobson@qut.edu.au

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

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