Biomass Conv. Bioref.DOI 10.1007/s13399-014-0123-9

ORIGINAL ARTICLE

Light olefins and transport fuels from biomass residues viasynthetic methanol: performance and cost analysis

Ilkka Hannula · Vesa Arpiainen

Received: 7 February 2014 / Revised: 15 April 2014 / Accepted: 16 April 2014© Springer-Verlag Berlin Heidelberg 2014

Abstract A thermochemical processing route frombiomass residues to light olefins (ethylene and propylene) isassessed by means of process simulation and cost analysis.A two-step process chain is proposed where (1) biomassresidues are first converted to synthetic methanol in a gasi-fication plant situated close to feedstock resources and (2)the produced methanol is transported to a steam crackingsite where it is further converted in a methanol to olefins(MTO) plant. Possibilities for heat and product integrationas well as equipment sharing with a steam cracking plantare discussed. Overall mass yields from dry biomass tolight olefins range from 169 to 203 kg/t. Based on cursorycapital cost estimates, the maximum methanol purchaseprice for such integrated MTO plants is estimated to be inthe range of 420–450 e/t.

Keywords Biomass · Gasification · Methanol · Olefins ·Biofuels

1 Introduction

Ethylene and propylene are the two largest volume petro-chemicals in the world. Their main applications are differentgeneral purpose plastics, like polyethylene and polypropy-lene, as well as elastomers and rubbers. Several techno-chemicals or chemical intermediates are also produced fromlight olefins as shown in Table 1.

Nearly all of the world’s supply of ethylene is pro-duced by steam cracking of hydrocarbon feedstocks such as

I. Hannula (�) · V. ArpiainenTechnical Research Centre of Finland,P.O. Box 1000, 02044 Espoo, VTT Finlande-mail: ilkka.hannula@vtt.fi

naphtha, ethane or LPG. About 55 % of the world’s propy-lene is produced as by-product from steam cracking ofpropane and higher hydrocarbons. Propylene is also sup-plied as by-product from fluid catalytic cracking (∼35 %)and propane dehydrogenation (∼5 %) [1]. Steam cracking isthe most energy-consuming process in the chemical indus-try, using globally approximately 8 % of the sector’s totalprimary energy use [2].

Rising prices of conventional petroleum feedstocks havedriven technology development to enable production ofolefins also from alternative feedstocks such as methanol.For example, four coal to olefins (CTO) projects haverecently been put into commercial operation is China1 andthe total production capacity of olefins in Chinese CTOprojects is approaching 3 Mt/a during the next 3 years [4].

However, synthetic products made from coal result innet greenhouse gas (GHG) emissions about double of thosefrom petroleum fuels [5]. These emissions could be loweredwith capture and storage (CCS) of the by-product CO2, butstill the net GHG emissions would be reduced only to lev-els comparable to those from petroleum fuels [6]. This isthe case also with CTO that uses synthetic methanol as anintermediate for olefin production.

European Union, in its effort to curb climate change anddecrease dependence on oil imports, has established a goalto introduce an increasing amount of renewable energy inthe transport sector. As a result, a great deal of RD&D workhas been focused to develop cost-effective gasification-based technologies for the production of synthetic fuelsfrom lignocellulosic biomass [7].

1These include Shenhua Baotou’s methanol to olefins (300 kt/a ethy-lene and 300 kt/a propylene) and Shenhua Ningmei’s methanol topropylene (500 kt/a propylene) [4].

Biomass Conv. Bioref.

Table 1 Global ethylene andpropylene end use markets in2011 [3]

Products from ethylene Ethylene end Products from Propylene end

use market propylene use market

% Mt % Mt

Polyethylene 60 76 Polypropylene 67 53

Ethylene oxide 14 18 Propylene oxide 8 6

Ethylene dichloride 11 14 Oxoalcohol 8 6

Ethyl benzene 7 9 Acrylonitrile 7 6

Others 8 10 Others 10 8

Total Sum 100 127 Total Sum 100 79

As these technologies are step by step gaining momen-tum and reaching maturity, the combination of syntheticbio-methanol manufacture with methanol to olefins technol-ogy would make it possible to produce light olefins fromlignocellulosic feedstocks. It is therefore of interest to eval-uate the potential role that this technology could have in thedecarbonisation of the petrochemical industry.

This paper presents an assessment of the co-productionof light olefins and transport fuels from lignocellulosicbiomass residues. We investigate (1) the performance of atwo-step process chain where methanol is first producedfrom forest residues via gasification followed by furtherconversion via methanol to olefins (MTO) technology, (2)two different end-uses for the by-product C4+: processingby an olefin cracker to boost the overall light olefin yieldor alkylation to produce renewable gasoline and (3) com-parative prospective economics for the two MTO processoptions. Detailed mass and energy balances are calculatedusing Aspen Plus� process simulation software. The overalleconomics are evaluated from the perspective of the MTOplant under varying by-product C4+ value, light olefin valueand total investment.

2 Process designs

2.1 Biomass to methanol

A simplified block diagram of a plant suitable for the con-version of biomass to synthetic methanol is shown in Fig. 1.The design of the plant is based on the author’s prior workdescribed in detail in [8]. The plant is operated with forestresidue chips whose properties (see Table 2) are represen-tative for typical by-products of pulp mills and mechanicalwood processing industries in Northern Europe.

The process begins by drying the feedstock from its ini-tial moisture of 50 to 15 wt% by an atmospheric bandconveyor dryer (belt dryer) operated with hot water (90 ◦Cin, 60 ◦C out). The dried biomass residue chips are thenpressurised with lock-hoppers and fed to a fluidised-bedsteam/O2-blown reactor operated at 5 bar and 850 ◦C. Thegasifier converts wood chips into raw product gas contain-ing CO, H2, CO2, H2O, CH4 and a small amount of higherhydrocarbons and tars. The raw product gas is processedby a downstream clean-up train employing hot-filtration at550 ◦C and catalytic reforming of hydrocarbons and tars to

OXYGENGASIFIER

BELTDRYER

ASU

AUXILIARYBOILER

Air N2

O2

H2S

Steam

Filter ash

Purge

Methanol

Steam

Recycle

Bypass

HOT-GASFILTER

OXYGENREFORMER

SOUR SHIFTWATER

SCRUBBER

CENTRIFUG.COMPR.

AGRMETHANOLCONVERTER

RECOVERY &DISTILLATION

CO2

Flue gas

Biomassresiduechips

SHAREDSTEAMCYCLE

Fig. 1 Simplified block diagram of a process capable of converting forest residue chips to synthetic methanol

Biomass Conv. Bioref.

Table 2 Properties of forest residue chips used as feedstock formethanol production [11]

Proximate analysis, wt% d.b.a

Fixed carbon 18.3

Volatile matter 80.6

Ash 1.1

Ultimate analysis, wt% d.b.

Ash 1.1

C 51.48

H 6.0

N 0.2

Cl 0.0

S 0.02

O (difference) 41.2

Other properties

HHV, MJ/kg 20.67

LHV, MJ/kg 19.34

Bulk density, kg d.b./m3b 293

Sintering temp. of ash, ◦C > 1000

aWt% d.b. weight percent dry basisb1 l batch, not shaken

light gases at 957 ◦C. The part of the model used to simulatethis three-step (gasification, filtration, catalytic reforming)process is validated with experimental data2 derived from an0.5 MWth process development unit (PDU) and described indetail in [9].

Following reforming, part of the gas is fed to an adia-batic reactor where water-gas shift reaction (1) is catalysedat around 300–400 ◦C to adjust gas module M = (H2-CO2)/(CO + CO2) to 2.03, a value required by the down-stream methanol synthesis [10]. After the shift reactor, thetreated gas is combined with by-pass and cooled down to40 ◦C while recovering heat for the plant steam cycle andthe belt dryer. The cooled gas is then compressed to 20 barto enable more efficient separation of acid gases by a phys-ical washing process using chilled methanol as solvent. Thetreated gas (now called synthesis gas or syngas) exiting theacid gas removal (AGR) unit is compressed further to 80bar, the operating pressure of the methanol converter.

CO +H2O = H2 + CO2 �H25 ◦C = −41.1 kJ/mol, (1)

Methanol is synthesised by hydrogenation of carbonoxide over catalysts based on copper oxide, zinc oxide orchromium oxide. All commercially available modern cat-alyst systems are based on Cu-ZnO-Al2O3 or Cr2O3 withdifferent additives and promoters. These catalysts allow theproduction of methanol at over 99.9 % selectivity with

2The PDU was run circa 4,000 h in pressurised oxygen-blown modeusing various wood residues as feedstock.

higher alcohols, ethers, esters, hydrocarbons and ketonesas primary by-products. Synthesis of methanol can bedescribed with the following reactions [12]

CO + 2H2 = CH3OH �H25◦C = −90.9 kJ/mol, (2)

CO2+3H2 = CH3OH+H2O �H25 ◦C = −50.1 kJ/mol.

(3)

Reaction heat is continuously removed from the reactorby boiling water to maintain essentially isothermic condi-tions at 260 ◦C and 80 bar. The unconverted synthesis gas isseparated from the reactor effluent, recompressed and recy-cled back to the feed side of the reactor until 95 % totalCO conversion is reached. The small amount of unconvertedgas (purge), which also contains methane left unconvertedin the reformer, is sent to the plant’s auxiliary boiler forcombustion.

In addition to small amounts of inorganic impurities, theproduced raw methanol contains also water that was formedas by-product of CO2 conversion. This water (along withsmall amounts of higher alcohols) might need to be sep-arated from methanol to achieve desired product quality.Higher purities can be achieved simply by adding more dis-tillation columns, thus contributing to additional capital andenergy costs. However, when methanol is produced exclu-sively for post-processing to olefins, a minimum distillationdesign is assumed to be a feasible approach.

Unconverted carbon from gasifier and unconverted syn-gas from methanol synthesis are combusted in an auxiliaryboiler to generate additional steam for the plant steam cycle.Syngas cooling before filtration and after reforming andshifting, together with methanol synthesis exotherm, countamong the other main sources for steam generation.

A small-scale extraction steam turbine is used to produceelectricity and to provide the plant with process steam at therequired pressure. The turbine can be implemented eitherwith a back pressure design to co-generate electricity anddistrict heat (DH), or with a condensing design to maximisethe output of electricity only. We have designed the plantsself-sufficient in terms of heat and steam, while electricityis balanced with the electric grid.

2.2 Methanol to olefins

Mobil (now ExxonMobil) discovered a zeolite-based ZSM-5 catalyst in the early 1970s that was capable of con-verting methanol to gasoline and olefins. This led to thedevelopment of methanol to olefins technology (MTO)in the mid-1980s as a spin-off to a methanol to gaso-line process demonstrated at that time in New Zealand.Later in the 1980s, scientists at Union Carbide developedSAPO-34 (silicoalumino-phosphates) catalysts that has high

Biomass Conv. Bioref.

selectivity for the MTO reaction [13]. The MTO develop-ment was transferred from Union Carbide to UOP in 1988but went largely unused until the mid-1990s when UOPteamed up with Norsk Hydro to build a pilot plant in Nor-way. A successful 100 bbl/day demonstration plant waslater operated in Germany with US and German governmentsupport [14]. Since then, Lurgi has developed its own ver-sion of the process, called methanol to propylene (MTP).The Chinese have also been active in this field and theDalian Institute of Chemical Physics has recently developeda similar process called DMTO [15].

Major differences between the MTP and MTO tech-nology are that the MTP process uses a ZSM-5 catalystsin a fixed-bed reactor to produce preferentially propylene,whereas the MTO technology is based on a MTO-100catalyst (a modified SAPO-34 catalyst) in a fluidised-bedreactor producing both ethylene and propylene with anadjustable product ratio. The analysis presented in this paperis based on the UOP/Hydro’s MTO technology.

A simplified block diagram in Fig. 2 illustrates a possibledesign for a stand-alone MTO process. The fluidised-bedreactor/regenerator system converts feed methanol into amixture of olefins, which is then fractionated to yieldpolymer-grade light olefins as major products.

The feed methanol is first compressed to 3 bar, preheatedand vaporised in heat exchange with reactor effluent andthen mixed with recycled methanol from the downstreamprocess. The methanol stream is then mixed with steam toincrease olefin selectivity and decrease catalyst deactivationin the reactor. The combined stream of methanol and wateris superheated to 310 ◦C and fed to a fast fluidised MTOreactor operating at 400–450 ◦C and 3 bar. At the presence

of a proprietary MTO-100 catalyst, a nearly complete(99.8 %) conversion of methanol is achieved with 80 %carbon selectivity to ethylene and propylene (see Table 3).Coke will gradually build-up on the catalyst surface and tomaintain activity, a portion of the catalyst is continuouslysent to a combustor (operating at 3.5 bar and 600 ◦C) wherethe coke is burned off with air before returning the regen-erated catalyst to the MTO reactor. The mass ratio betweenethylene and propylene in the effluent (stream 2 in Fig. 2)can be varied from 0.75 to 1.5 by adjusting the operat-ing severity. Higher temperature will lead to more ethylenebeing produced [16], although highest overall yield to lightolefins (ethylene plus propylene) is achieved with aboutequal amounts of both [17].

The reactor effluent is cooled down to 240 ◦C in afeed/effluent heat exchanger and then further to condensthe water and unconverted methanol by a scrubber (labelledquench in the flowsheet). The recovered methanol is recy-cled back to the reactor. The bottom stream of the strippercontains most of the water contained in the MTO reactor’seffluent and is sent to waste water treatment after exchang-ing heat with reactor feed preheater. The gaseous effluent iscompressed to 25 bar and flashed at 33 ◦C in a phase separa-tor to produce a vapour stream and a condensate stream withtwo liquid phases. The aqueous phase is separated from thecondensate and sent to the stripper while the organic layeris stripped in a separate column and the produced organicconcentrate is sent to a downstream depropaniser (labelledDe-C3).

Acid gases from the phase separator’s vapour stream areremoved by caustic wash. The treated acid-free effluent isthen cooled to 22 ◦C, dried with a molecular sieve, cooled

REACTOR &

REGENERATOR

C2H2

REACTOR

DE-C2

DRYER

CAUSTIC

WASH

QUENCH

DE-C1

C2 SPLITTER

DE-C3

C3 SPLITTER

Crude

methanol

Air

PHASE

SEPARATOR

COMPRESSOR

CONDENSATE

STRIPPER

COMPRESSORFluegas

Waste

water

Fuelgas

Ethylene

Ethane

Propylene

Propane

WATER

STRIPPER

1

25

6

78

9

3

4

10

BLOWER

OLEFIN

CRACKING

C4+C4+ residue

C2/C3

11

PUMP

Steam

Fig. 2 Simplified block diagram of a methanol to olefins plant modelled in this work. Simulation results for the numbered process streams aregiven in Table 9

Biomass Conv. Bioref.

further to 10 ◦C and sent to a de-ethaniser (De-C2) wheremajority of ethylene is recovered overhead and most of thepropylene from the bottom (condenser temperature −25 ◦C,reboiler 66 ◦C). The overhead vapour is compressed to 33bar and sent through an acetylene converter (C2H2 reactor)where the small amount of acetylene produced in the MTOreactor is hydrogenated to ethane over a palladium-basedcatalyst. The treated effluent is then chilled down to −20 ◦Cand fed to a demethaniser (De-C1) that produces methane-rich fuelgas overhead (5) and a mixture of C2 hydrocarbonsfrom the bottom. The fuelgas is routed through a pressureswing adsorption unit that recovers 86 % of the hydrogencontained in the stream. After hydrogen recovery, the restof the gas is directed to combustion. Very low temperatures(−90 ◦C in the condenser) are needed to carry out this sep-aration. The C2 stream from the bottom is directed to aC2-splitter column that produces a polymer-grade ethylenestream overhead (6) and an ethane-rich (about 70 mol%)by-product stream from the bottom (7). The bottom streamfrom the de-ethaniser (De-C2) is mixed with the bottomsfrom the organic layer stripper (condensate stripper) andsent to a depropaniser (De-C3). The overhead stream goesto a large C3-splitter producing polymer-grade propylene(8) overhead and a propane-rich (around 60 mol%) by-product (9) from the bottom. The De-C3 bottoms (stream10) consists heavy hydrocarbons characterized as a C4+stream.

2.3 Olefin cracking process

Two different MTO plant designs are examined in thispaper. They differ from each other in the way the by-product C4+ stream is processed. In our ‘base case MTO’design, the C4+ by-product is sent to alkylation unit, where1-butene and 2-butene react with isobutane to form valu-able high-octane alkylates used as gasoline additives. In our‘advanced MTO’ design, the C4+ stream is sent to an olefincracking process (OCP) where it is converted to higher valuepropylene and ethylene.

The advanced MTO process can reach close to 90 %overall carbon selectivity to ethylene and propylene frommethanol, a marked improvement from the 80 % of the basecase MTO. While the propylene to ethylene (P/E) ratio ofconventional MTO is about 1, in OCP the P/E ratio is 3.5–4.0. The P/E of the combined MTO + OCP process canthus range from 1.3 to 2.1 [13, 19, 20]. The integration ofMTO with OCP is fairly straightforward because the recov-ery section of the MTO unit remains unchanged, needingonly to be resized to accommodate the added circulation toand from the olefin cracking process [21].

The olefin cracking process is developed by Total Petro-chemicals and comprises of selective hydrogenation reac-tor, olefin cracking reactor and two fractionating columns

(depropaniser and debutaniser). In the selective hydrogena-tion reactor, diolefins and acetylenes present in the feedare converted to mono-olefins to prevent their conversionto coke and further to methane later in the process. Theselective hydrogenation is performed at relatively mild con-ditions (30–200 ◦C, 4.5–22 bar) in liquid phase using acylindrical fixed bed reactor with alumina catalyst [22].

In the olefin cracking reactor, heavy C4+ olefins arecracked down to light olefins in the C2 to C3 range undergaseous phase conditions and in the presence of an olefincracking catalyst [22]. The reactor is operated at 500–600 ◦C and 1–5 bar [13]. For the assumed mass yieldstructure of the olefin cracker process, see Table 3.

The light olefin product stream is recovered from over-head of the depropaniser fractionating column. The debu-taniser fractionating column is used to collect and redirecta portion of the depropaniser bottoms stream for processrecycle and also to remove process purge comprising C4

and heavier hydrocarbons to avoid the build-up of paraffiniccompounds. The depropaniser and debutaniser columnsoperate at 8 to 21 bar [22].

3 Simulation methodology and results

3.1 Biomass to methanol

Mass and energy balances are simulated for a methanolplant processing 100 MWth (LHV) of wet biomass. Theresults are calculated separately for two different plantdesigns employing either (a) condensing, or (b) combinedheat and power (CHP) steam system design. When operated

Table 3 Mass yield structures used for the simulation of the MTO andolefin cracking process [13, 16, 18]

Component MTO OCP

H2 0.0648

CO 0.0128

CO2 0.0519

CH4 1.0486 2

C2H4 16.7210 19

C2H6 0.3581 2

C3H6 16.7208 71

C3H8 0.2245 4

C4H8 5.1564 2

C4H10 0.0389

C5H10 2.1076

Coke 1.3255

H2O 56.1692

SUM 100.0000 100

Biomass Conv. Bioref.

with condensing turbine, the steam system produces11.0 MW of electricity leading to −1.5 MW deficit aftersatisfying the plant’s −12.5 MW parasitic electricity con-sumption. With back pressure turbine the electricity outputdecreases to 8.3 MW (−4.2 MW net electricity output) butis compensated with 12.3 MW of co-product heat that canbe utilised as district heat or low pressure industrial steam.Breakdown of the simulated electricity and steam balancesis given in Table 4.

Main simulation results for the methanol synthesis aregiven in Table 5. From the 100 MWth of wet biomass fedinto the process, 77 MWth of ultra-clean synthesis gas isgenerated. In the methanol loop, the synthesis gas is fur-ther converted to 60 MW of methanol, while the reactionexotherm is used to generate 3.5 kg/s of saturated steam at43 bar, which is then used to satisfy part of the methanolplant’s own steam consumption. And 4.9 MW worth of gasis continuously purged from the synthesis to avoid build-upof inerts in the recycle loop.

Fuel production is unaffected by the choice of steamcycle design. The overall thermal efficiency η is calculatedaccording to the following equation

η = mMeOH ×HMeOH +QDH

mbiom ×Hbiom

, (4)

Table 4 Process simulation electricity and steam balances for abiomass to methanol plant processing 100 MWth (LHV) of wetbiomass (or 511 t/day dry basis)

Electricity balance, MW

On-site consumption −12.4

Oxygen production −3.6

Oxygen compression −0.7

Feeding and lock-hopper −0.2

Feed drying −0.7

Syngas compression −5.2

Acid gas removal −0.9

Synthesis −0.3

Blowers + pumps −0.3

Miscellaneous −0.6

Gross production

With condensing turbine 11.0

With back pressure turbine 8.3

Steam balance, kg/s

On-site consumption 8.0

Gasifier 2.4

Reformer 1.4

AGR solvent regeneration 1.1

Deaerator 1.4

Economiser 1.7

Gross production 16.1

Table 5 Process simulation results for the methanol synthesis

Methanol synthesis

Synthesis gas input MW 77

Methane in syngas MW 1.1

Water at reactor inlet mol% 0.0

Inlet pressure to synthesis bar 80

Reactor inlet temp. C 260

Reactor outlet temp. C 260

Per-pass CO conversion % 50.8

Total CO conversion % 95.4

Reactor steam output kg/s 3.5

Unconverted gas to boiler MW 4.9

where m denotes massflow (kg/s), H heat of combustion3

(MJ/kg) and QDH possible district heat co-product (MW).For the discussed methanol plant designs, the efficiencyfrom wet biomass to methanol is 60 % (LHV), which cor-responds to a mass yield of 0.5108 kg/kgbiomass,dry. Whendistrict heat is produced along with fuel, the overall effi-ciency from biomass to methanol and DH increases to 72 %.Key performance parameters for the methanol plant aresummarised in Table 6.

3.2 MTO and OCP

Simulation results for the MTO and OCP are continuationof the author’s prior work reported in [23, 24]. The MTOreaction is simulated with an Aspen’s RYield block assum-ing detailed product yield structure based on Kuechler andLattner [18] and Wan [16]. The yield structure used to simu-late the olefin cracking process is less detailed and deducedfrom [13]. Both yields are given in Table 3.

The energy consumption of the olefin cracker is cal-culated using Aspen’s RGibbs block. When only ethyleneand propylene are assigned as possible products, the targetpropylene/ethylene (P/E) ratio of 3.7 is achieved at 590 ◦Cand 191 kJ/kgMeOH of heat (2.6 MJ/kgC4+) is required torun the endothermic reaction. In our advanced MTO plantdesign, this heat is provided by combusting a fraction of theby-product fuelgas from the main MTO process.

Table 9 presents simulation data on selected streams (seeFig. 2 for numbering) of the MTO process. All the massflows are presented in relation to methanol input. Streams6 (ethylene) and 8 (propylene) are the main products bothhaving ≥ 99.5 % purity. By-products include ethane-rich(∼70 mol%) stream 7, propane-rich (∼60 mol-%) stream9 and methane-rich (∼60 mol%) stream 5. Stream 10 con-tains the C4+ fraction of the effluent that in the base case

3The heat of combustion is defined on lower heating value basis andcalculated for the wet biomass prior drying.

Biomass Conv. Bioref.

Table 6 Key performance results for the simulated biomass tomethanol plant designs

Condensing CHP

Thermal efficiencies (LHV)

MeOH/wet biomass 60 % 60 %

DH out/wet biomass 12 %

MeOH + DH/wet biomass 72 %

Utilities, kJ/kgbiomass

Net electricity to grid −227 −678

District heat (90 C) 0.0 2084

Mass yield, kg/kgbiomass

MeOH/biomass (dry) 0.5108 0.5108

MTO design is sent to alkylation, but in the advanced MTO(MTO + OCP) design is recycled back to the MTO via olefincracking process to boost the overall light olefin yield frommethanol.

3.3 Refrigeration

Fractionation of the MTO effluent requires cooling belowambient temperature, which is provided by compressionrefrigeration system. The required cooling duties are simu-lated and the related refrigeration work requirements esti-mated using relation

COPCarnot = T evap

(T evap − T cond)= QC

W, (5)

where COPCarnot is the ideal coefficient of performance,Tevap is evaporation temperature (K), Tcond is condensationtemperature (K), QC is cooling duty and W is refrigerationwork. The actual performance of an refrigeration cycle isassumed to be 0.6 times the ideal [25]. Simulated coolingduties and related work requirements for both plant designsare summarised in Table 7. For both designs, the largestcooling duties are required in the C2-splitter’s condenser(51 and 58 kJ/kgMeOH ), while C1-column’s condenserhas the largest refrigeration work consumption (29 and31 kJ/kgMeOH ), mainly due to the low (−90 ◦C) evapora-tion temperature that leads to low COP for the refrigerationcycle.

The required reboiler duties are also simulated. They arecovered with 15 bar steam that is assumed to be availablefrom the steam grid of the steam cracker’s site. Accordingto the results, the reboilers’ combined steam requirement is0.17 kg/kgMeOH for the MTO and 0.24 kg/kgMeOH for theadvanced MTO design.

3.4 Products and utilities

Table 8 summarises the main simulation results for themethanol to olefins process. In the MTO design, the C4+

stream is sent to alkylation, while in the advanced MTOdesign it is sent to olefin cracking process and the C1–C3

fraction of the OCP effluent (see Table 3 and Fig. 2) is recy-cled back to the MTO process. In the MTO design, equalamounts of ethylene and propylene are produced on massbasis, while in the advanced MTO design the simulated P/Eratio is 1.2.

The overall mass yields from dry biomass to light olefinsare 0.169 and 0.203 kg/kg for the MTO and advanced MTO,respectively. The two plant designs also consume differ-ent amount of utilities. In both cases, the largest consumerof electricity is MTO effluent’s compression that requires106 kJ of electricity for every kilogram of methanol fedinto the process. The refrigeration is the second largestat 62 kJ/kgMeOH with the advanced MTO design requir-ing additional 7 kJ/kgMeOH to offset the increased coolingduties caused by fractionation of additional light olefinsfrom the olefin cracker.

Steam is both produced and consumed in the MTO pro-cess. In the conversion area, heat is recovered both fromthe exothermic MTO reaction and from the regenerator’sfluegas to produce 0.33 kg/kgMeOH of high-pressure super-heated (125 bar, 520 ◦C) steam, while 0.32 kg/kgMeOH

of saturated low pressure (15 bar, 198 ◦C) steam is usedto vaporise water prior mixing with the methanol feed.In the separation area, 0.17 kg/kgMeOH of low pressuresteam is required to heat the reboilers of the MTO and0.24 kg/kgMeOH of the advanced MTO design. Whencombined, these amount to 0.49 and 0.56 kg/kgMeOH low-pressure steam requirement for the MTO and advancedMTO designs, respectively.

4 Cost analysis

4.1 Integration of MTO with existing steam cracking plant

A fully functional MTO process requires all the main equip-ment illustrated in Fig. 2 plus equipment to provide utilitieslike steam and refrigeration. However, major savings incapital investment could possibly be attained by couplingthe MTO process with an existing steam cracking plant.This way the MTO process could benefit from the steamcracker’s fractionation capacity and utility equipment, thuslimiting the required investment only to the methanol con-version section, i.e. the grey shaded boxes shown in Fig. 2.Such an ‘integrated MTO’ design would produce two inter-mediate streams: a C1–C3 and a C4+ stream (indicated withnumbers 3 and 4 in Fig. 2) that would be integrated to exter-nal units for further processing. In our designs, the C1–C3

stream is routed to the fractionation part of the steam crack-ing plant and the C4+ stream either to alkylation or olefincracking depending on the examined case.

Biomass Conv. Bioref.

Table 7 Simulated cooling duties (QC) and corresponding work (W) requirements of the compression refrigeration system

COPCarnot MTO, kJ/kgMeOH MTO+OCP, kJ/kgMeOH

QC W QC W

De-C2 precooler 18.9 38 3 47 4

De-C2 column 4.9 32 11 36 12

De-C1 precooler 5.6 27 8 31 9

De-C1 column 1.6 28 28 30 30

C2-splitter 7.1 51 12 58 14

SUM 176 62 201 69

The results are calculated separately for the MTO and advanced MTO (MTO + OCP) plant designs. The actual COP is estimated to be 0.6 timesthe ideal. All duties are given in kJ/kgMeOH

To make sure that there is enough capacity availableat the steam cracking plant for the fractionation of theMTO’s C1–C3 stream, we assume that a steam crackingoven of comparable size is taken off-line before start-upof the MTO process. Due to differences in yield struc-tures between steam cracking and MTO, equal amounts ofethylene produced is not an ideal indicator for quantify-ing equal capacities. For this reason, we have chosen tocompare capacities based on equal amount of light olefins(ethylene + propylene) produced. We assume 0.57 lightolefin mass yield for naphtha steam cracking and 0.33 (seeTable 9) for MTO. Based on these assumptions, 1.73 kg ofmethanol is required to replace 1 kg of naphtha to producecommensurate amount of light olefins.

We assume that the process energy demand in naphthasteam cracking is met by combustion of by-products andheat recovery from flue gases and waste heat. We furtherassume that all the required steam is generated within theprocess leading to zero net steam import or export. As aresult, we expect that the replacement of a naphtha ovenwith an MTO reactor will not cause major changes in theheat and steam balance of the entire steam cracking plant.

4.2 Capital cost estimates

Based on discussions with industry experts, we assume thata modern naphtha cracking oven with a feed input of 50 t/hhas total capital investment in 2014 (TCI) of 45 Me. Wefurther assume that an MTO conversion section of commen-surate size (86 t/h MeOH) requires double the TCI at 90Me4.

We then want to estimate the MTO’s capital investmentat a scale that would fit better with the limited nature of

4Higher costs are due to more expensive double fluidised-bed reac-tor/regenerator system, reactor internals, catalyst and combustion airblower.

biomass resources. We set the target size of the MTO to30 t/h of methanol consumption and scale the TCI with thefollowing relation

C = C0 ∗(S

S0

)0.85

, (6)

where So and Co denote the scale and cost of referenceequipment and S and C the scale and cost of equipmentat the target scale. The cost downscaling exponent of 0.85for the methanol conversion section is suggested by Wan

Table 8 Key performance results for MTO and advanced MTO(MTO + OCP) plants based on Aspen Plus� simulation. Yields referto final yields after separation and recycling

Product streams, kg/kgMeOH MTO MTO+OCP

Ethylene 0.1659 0.1798

Propylene 0.1655 0.2173

C4+ 0.0737 0.0015

H2 0.0006 0.0006

Fuelgas* 0.0114 0.0080

Ethane-rich 0.0049 0.0067

Propane-rich 0.0025 0.0054

Utilities MTO MTO+OCP

Work, kJ/kgMeOH

Air blower (MTO) −35 −35

Compression (MTO) −106 −106

Compression (OCP) −33

Cryogenic work (MTO) −62 −69

Steam, kg/kgMeOH

HP steam (MTO) 0.33 0.33

LP steam (MTO) −0.49 −0.56

Overall mass yields, kg/kg

Ethylene/biomass (dry) 0.0847 0.0918

Light olefins/biomass (dry) 0.1693 0.2028

*After H2 separation by PSA and combustion in OCP

Biomass Conv. Bioref.

Table 9 Simulation results for major MTO process streams as numbered in Fig. 2

Stream 1 2 3 4 5 6 7 8 9 10

Pressure, bar 3 3 25 24 31 31 31 21 21 24

Temperature, ◦C 310 410 33 50 −90 −12 10 50 58 128

Flow, kg/kgMeOH 1.225 1.212 0.411 0.013 0.012 0.166 0.005 0.166 0.003 0.074

Compositions, mol%

H2 0.006 0.027 0.320

CO 0.003

CO2

CH4 0.012 0.054 0.648 0.001

C2H4 0.106 0.494 0.030 0.998 0.189

C2H6 0.002 0.010 0.002 0.684 0.001

C3H6 0.071 0.329 0.008 0.127 0.995 0.347 0.016

C3H8 0.001 0.004 0.001 0.004 0.571 0.001

C4H8 0.016 0.057 0.983 0.082 0.736

C4H10 0.009 0.005

C5H10 0.005 0.025 0.242

CH4O 0.714

H2O 0.286 0.780

SUM 1.000 1.000 1.000 1.000 1.000 1.000 1.000 1.000 1.000 1.000

Results for the advanced MTO design are not included in this table. Compositions are given in mole fractions while residual concentrations smallerthan 0.1 % are omitted for the sake of readability

[16]. Based on this relation, we estimate the TCI of an MTOconversion section using 30 t/h of methanol to be 37 Me.

Based on our simulation results (see Table 9), the C4+by-product flow from 30 t/h methanol MTO plant is 2.2 t/h.We assume that the cost of MTO conversion section and theolefin cracking process of comparable size (as measured interms of feed input) are equal and by using again the above-described relation, we estimate the TCI of such OCP tobe 4 Me.

4.3 Prices and methodology

Financial parameters used in the cost analysis are summa-rised in Table 10. The annual capital charges are calculatedfrom TCI using 0.12 annuity factor, which corresponds with10 % interest and 20 years lifetime. The yearly operatingand maintenance (O&M) costs5 are assumed to be 4 % ofthe TCI [26] and the 7,890 annual operating hours corre-spond with 0.9 on-stream factor. Ethylene is valued at 1,100e/t and propylene at 1,200 e. The C4+ stream is valuedeither at 600 e/t or 1,015 e/t depending on whether theend-product is considered as conventional motor fuel or bio-fuel6, hydrogen is valued at 1,500 e/t and the rest of the

5Following breakdown is assumed for the O&M: personnel costs0.5 %, maintenance & insurances 2.5 %, catalysts & chemicals 1 %.6The C4+ is valued at 80 % of the end-product value, which we assumeto be 750 e/t for conventional motor fuel and 1,269 e/t for biofuel

by-products (fuel gas, ethane and propane) at 500 e/t. Costsrelated to buying or selling utilities are valued at 50 e/MWhfor electricity, 35 e/t for high pressure steam and 30 e forlow pressure steam.

Table 10 Financial parameters assumed for the purpose of costanalysis

Investment factors

Annuity factor (10 %, 20 a) 0.12

Annual O&M cost factor 0.04a

Annual operating hours 7890

Investment support, Me 0

Product values, e/tonne

Ethylene 1,100

Propylene 1,200

H2 1,500

C4+ 600/1,015b

Fuelgas 500

Ethane 500

Propane 500

Utilities

Electricity, e/MWh 50

High pressure steam, e/t 35

Low pressure steam, e/t 30

aFraction of Total Capital InvestmentbWhen end-product conv. motor fuel / biofuel

Biomass Conv. Bioref.

Capital cost estimates provide the basis for evaluatingthe prospective economics of methanol to olefins. We dothis by using maximum methanol purchase price (MMPP)as the economic indicator, and calculate it according to thefollowing equation:

MMPP(e/t) = P + U + C +O

M, (7)

where P is revenue from selling the hydrocarbon products,U is the cost of utilities (net electricity, high and low pres-sure steam), C is annual capital charges and O is operatingand maintenance costs. Incomes are considered positive andexpenses negative costs. The sum of these annual costs (e/a)is divided by M , which is the annual methanol input (t/a) tothe MTO process. Defined in this way, the MMPP (e/t) indi-cates the maximum price that can be paid from the MTO’sfeedstock.

4.4 Methanol selling price estimates

Methanol is a global product that is mainly traded byquarterly contract. In Europe, the contract price of fossilmethanol for 4Q/2013 was 408 e/t FOB Rotterdam.

In contrast, the worlds bio-methanol market is extremelysmall and price information is not easily available. However,several production cost estimates for synthetic bio-methanolare available in the public domain [8, 27–32]. Talmadgeet al. [32] summarise minimum selling price estimates fora variety of synthetic biofuels and chemicals based oninputs from selected literature sources. For bio-methanol,their minimum selling price estimates are 295 ± 67 e/tfor mature (Nth plant) technology and 469 ± 95 e/t for apioneer plant7. Based on these figures, it appears that theproduction costs of the first bio-methanol plants will beclearly higher than current market prices, although in thelong term, and under favourable conditions, synthetic bio-methanol appears to have the potential even to undercut thecurrent market prices of fossil methanol.

4.5 Sensitivity analysis

We now proceed to evaluate the prospective economics ofmethanol to olefins in the form of a sensitivity analysis. Weuse maximum methanol purchase price (MMPP) as the eco-nomic indicator, which represents the maximum price that aplant is able to pay for its feedstock. We examine the MMPPas a function of by-product C4+ value, bio-olefin premiumand total investment.

7Both plants having 10 % internal rate on return and paying 12e/MWh for the feedstock

4.5.1 Impact of C4+′s value

Figure 3 illustrates the impact of C4+′s value to the maxi-mum methanol purchase price. In alkylation, the C4+ streamis converted to premium gasoline blending stock. When bio-methanol is used as feedstock for the MTO, blending stockproduced from the C4+ becomes a biofuel component. Var-ious mandates and obligations are already in place in thebiofuels market, making it possible to charge price premi-ums. We price conventional petroleum-derived motor fuelat 750 e/t (before taxes at the refinery gate) and biofuelat 1,269 e/t(gasoline eq.), which was the FOB price forBrazilian T2 ethanol in Rotterdam on June 2013 [33].

According to the results, if the seller is unable to chargeprice premium on the by-product alkylate, the maximummethanol purchase price is 405 and 434e/t for the MTO andadvanced MTO designs, respectively. However, when man-dates are in place, MMPP for the MTO design rises to 435e/t and becomes on a par with advanced MTO. This hap-pens because only a very small amount of C4+ is producedin the advanced MTO design (most of it is cracked to lightolefins in OCP), making its economics non-sensitive to thevalue of alkylate.

4.5.2 Impact of bio-olefin premium

Figure 4 illustrates the impact of bio-olefin premium to themaximum methanol purchase price. If the olefin seller isunable to charge any premium from bio-olefins relative tothe prevailing market price (ethylene 1,100 e/t and propy-lene 1,200 e/t), the MMPP is 435 e/t for both MTO andadvanced MTO. However, if an bio-olefin premium exists,

405

435434 435

300320340360380400420440460480500

600 1015

MM

PP

, € /t

onne

C4+ value, € /tonne

MTO MTO+OCP

Fig. 3 Maximum methanol purchase price calculated with two differ-ent values for the C4+: 600 e/t (when end-product conventional motorfuel), and 1,015 e/t (when end-product biofuel)

Biomass Conv. Bioref.

420430440450460470480490500510520

0 50 100 150 200

MM

PP

, €/to

nne

Bio-olefin premium, €/tonne

MTO+OCP

MTO

Fig. 4 Maximum methanol purchase price calculated as a function ofbio-olefin premium

then every 100 e/t (9 %) increase in the price of light olefinswill lead, on average, to 36 e/t increase in the MMPP. Forexample, at a 200 e/t premium, the MMPPs is 501 e/t and514 e/t for base case MTO and advanced MTO, respec-tively. The advanced MTO design is more sensitive to theolefin value because of the increased output caused by theOCP.

4.5.3 Impact of total capital investment

We further examine the impact of a change in the total cap-ital investment (TCI) to the MMPP. Figure 5 illustrates theimpact of a change in total capital investment (TCI) to theMMPP. If the TCI would be 50 % less than in our estimate,then the MMPP would increase, on average, by 13 e/t to448 e/t for the MTO and advanced MTO. A 50 % increasein the TCI would have a similar size impact to the otherdirection, lowering the MMPP to 423 for MTO and 421 e/t

420

425

430

435

440

445

450

-50 % -25 % 0 % 25 % 50 %

MM

PP

, €/to

nne

Total Capital Investment, change

MTO+OCP

MTO

Fig. 5 Maximum methanol purchase price calculated as a function ofchange in total capital investment

for advanced MTO. The advanced MTO design is slightlymore sensitive to changes in TCI due to the higher abso-lute investment caused by the addition of olefin crackingprocess.

5 Summary and conclusions

Technology for the catalytic conversion of methanol toolefins is commercially ready, with several commercialprojects already operating and more in planning stages inChina. In addition, most of the components needed forthe conversion of biomass to methanol are commerciallymature, making near-term deployment of such plants pos-sible. When these technologies are combined, they allowthe production of light olefins from lignocellulosic biomass.Based on our analysis, 169–203 kg of light olefins couldbe produced from 1 t of dry biomass with such combinedprocess.

Two different MTO plant designs, differing in the waythe by-product C4+ stream from the main process istreated, have been examined. In the ‘base case MTO’design the C4+ stream is alkylated to high-octane blend-ing stock and in the ‘advanced MTO’ design it is crackedto produce more high value light olefins. The light olefinyields are 0.331 kg/kgMeOH for the base case MTOand 0.397 kg/kgMeOH for the advanced MTO design.For the overall production chain, the mass yields are0.1693 kg/kgbiomass,dry and 0.2028 kg/kgbiomass,dry for thebase case MTO and advanced MTO, respectively.

Based on our prospective cost analysis, under theassumption that the by-product blending stock (producedfrom the C4+ stream) is valued as a biofuel component, themaximum methanol purchase price (MMPP) is calculatedto be 435 for both MTO designs. This happens because,although the advanced MTO has a higher light olefins out-put than base case MTO, it produces less C4+ that is alsohighly valuable under the biofuel assumption. If the olefinseller is able to charge premium from the bio-olefins, thenevery 100 e/t (9 %) increase in the price of light olefins willlead, on average, to 36 e/t increase in the MMPP. In addi-tion, a 50 % change in the capital investment will cause,on average, a 13 e/t change in the MMPP. These resultsare based on the underlying assumption that major savingsin capital investment can be attained through integrationbetween the MTO and an existing steam cracking plant.

Acknowledgments The authors thank Lotta Sorsamaki, Ali Harlinand Kai Sipila (VTT) for helpful discussions during the preparation ofthis paper, Tom Kreutz (Princeton University) for helpful reviews of anearly MTO simulation model and Eric Larson (Princeton University)for assistance with early cost analysis. The research leading to theseresults was funded by VTT Technical Research Centre of Finland.

Biomass Conv. Bioref.

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