Technology Profiles (Chemical Engineering)

8/16/2019 Technology Profiles (Chemical Engineering)

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D,L-methionine production via the

carbonate process

M

ethionine is an essential amino acidthat is not synthesized by animals, so

it must be obtained from their diets.Synthetic methionine is traditionally commer-cially available in two varieties: D,L-methionine(DLM), a racemic mixture of the two stereoiso-mers, and methionine hydroxy analog (MHA).For animal nutrition, the two are equivalent.Globally, 98% of methionine produced is usedas an animal feed additive, especially in thepoultry market. Methionine is also used inpharmaceuticals and cosmetics.

The processIn the process shown in Figure 1, DLM isobtained via the carbonate process from3-methylthiopropionaldehyde (MMP) and

hydrogen cyanide (HCN), without forming co-products. The process presented here is similarto one developed by Evonik Industries AG (Es-sen, Germany; www.evonik.com), and can bedivided into four main areas: hydantoin produc-tion; hydantoin hydrolysis; DLM production andfiltration; and drying. Hydantoin correspondsto the intermediate 5-(2-methylmercaptoethyl)-hydantoin generated in the process.Hydantoin production. MMP, HCN and anaqueous solution of carbon dioxide (CO2) andammonia (NH3) are fed to the hydantoin prima-ry reactor. A gas stream containing CO2, NH3 and impurities leaves the reactor and is washedwith MMP and water in a scrubbing column.

Purified CO2 is released in the overheads andused in DLM production. The reaction mixtureis fed to a hydantoin secondary reactor, wherethe reaction is completed.Hydantoin hydrolysis. The hydantoin productstream and potassium carbonate (K2CO3)are fed to the hydrolysis reactive column,which contains zirconium fittings that serveas catalysts and prevent corrosion. CO2 andNH3 are liberated in the overhead streamof the column and recycled to the hydantoinprimary reactor. The bottom stream, contain-ing potassium methioninate (KMET), is sent tothe DLM production portion of the process.K2CO3 is obtained from the regeneration of a

potassium bicarbonate (KHCO3) solution fromDLM filtration. This occurs in a decarbonator,where CO2 is separated in the overheads anddirected to the DLM reactor.DLM production and filtration. The KMET streamand recycled CO2 are fed to the DLM reactor,forming KHCO3 and DLM. The methionine isfed to a vacuum crystallizer and the crystals arevacuum-filtered. The resulting KHCO3-contain-ing filtrate is sent to the decarbonator.Drying. The methionine cake is directed to a

dryer, where water is evaporated and sent to awash column before being used in methioninewashing during filtration. Methionine product(99 wt.%) is sent to packaging and storage.

Economic performanceAn economic evaluation of the process wasconducted, assuming a 150,000 ton/yr methi-onine unit erected on the U.S. Gulf Coast andintegrated with nearby MMP and HCN facili-ties. The storage autonomy is equal to 30 daysof operation for methionine, aqueous NH3 andliquefied CO2. The estimated capital invest-ment is about $390 million, while operatingexpenses are about $2,600 per ton of product.

Global perspectiveIn the past few years, methionine globaldemand has increased at an average rate of5%/yr, and is expected to increase further.This expansion of methionine demand is drivenby rising meat consumption and, in particu-lar, growth in global poultry, associated withpopulation growth and increased purchasingpower, mainly in developing countries. Also,there has been consolidation in the agriculturalsector, leading to larger, more industrialized

farms with more efficient animal nutrition withhigher inclusion rates of amino acids.In this context, Asian markets present strong

methionine demand, driving its manufacturetoward that region. This can be observedthrough the methionine capacity expansionplanned in this area (Figure 2). n

Editor’s Note: The content for this column is suppliedby Intratec Solutions LLC (Houston; www.intratec.us)and edited by Chemical Engineering. The analysesand models presented are prepared on the basis ofpublicly available and non-confidential information.The content represents the opinions of Intrateconly. More information about the methodology forpreparing analysis can be found, along with termsof use, at www.intratec.us/che.

1 ) Hydantoin primary reactor

2 ) Hydantoin secondary reactor

3 ) Scrubbing column

4 ) Hydrolysis reactive column

5 ) Decarbonator

6 ) DLM reactor

7 ) Vacuum crystallizer

8 ) Vacuum filter

9) Dryer

10) Vacumm wash column

CW Cooling water

ST Steam

PW Process water

11) Cooling tower

12) Steam boiler

3

5

78

2

9

10

CW

ST

MMP

MMP+

Impurities

MMP

HCN

CO2 make-up

H2O

NH3 make-up

CO2 make-up

K2CO3 make-up

ST

CW

CW

CW

ST

Methionine topackaging and storage

PW

Toflare

1

4

6

11

12

FIGURE 1. D,L-methionine production via carbonate process

FIGURE 2. Worldwide methionine production capacity

4 0 0

5 8 0

3 0 0

9 0

8 0 1 5 0

n Total current capacity (2014) = 1,360,000 ton/yr

n Total planned capacity (2016) = 240,000 ton/yr



Extracting 1,3-Butadiene

from a C4 Stream

The organic compound 1,3-butadieneis a petrochemical commodity used asa raw material for the production of

rubbers and plastics, such as polybutadienerubber (PR), styrene butadiene rubber (SBR),and acrylonitrile butadiene styrene (ABS).These materials are mainly applied in themanufacture of automotive parts, tires andcables. Also, 1,3-butadiene is used as anintermediate in the manufacture of severalchemicals, such as adiponitrile, the rawmaterial for nylon production.

1,3-Butadiene is generally recovered fromC4 streams that are generated as byprod-ucts of ethylene manufacture by naphtha-based steam cracking. These C4 mixturesare composed mostly of butadiene andbutenes, with smaller amounts of butanes

and acetylenes. To obtain the 1,3-butadieneproduct stream, it must be extracted from theC4 mixture. This can be accomplished usingseveral technologies, such as separation bysolvent extraction.

The processThe process depicted in Figure 1 is similarto BASF’s (Ludwigshafen, Germany; www.basf.com) process for butadiene extrac-tion using aqueous N-methylpyrrolidinone(NMP) as a solvent.Extractive distillation. The C4 mixturecontaining butadiene, butenes, butanesand acetylenes is fed to the first extrac-tive column, where recycled NMP solventis added to the top section. The columnoverhead stream, which consists of butanesand butenes (raffinate-1), is sent to storage.The bottom stream, containing butadiene,acetylenes and some butenes absorbed inNMP, is fed to the top of the rectifier col-umn. In this column, remaining butenes areseparated in the top stream and recycledto the first extractive column, while thebottoms stream, containing mainly NMP,is sent to the degassing section. Butadieneand some acetylenes are separated as a

side stream of the rectifier column, which

is fed to the second extractive column.There, recycled NMP solvent is added tothe top to absorb the acetylenes, whichare recovered in the bottoms and returnedto the rectifier column. Crude butadiene isrecovered in the overheads and sent to thebutadiene-purification section.Butadiene purification. In this section, thecrude butadiene stream is fed to the pro-pyne column, where propyne is separatedin the overheads. The bottoms stream ofthis column is sent to the butadiene distilla-tion column, where 1,3-butadiene productis obtained in the overhead stream andheavy ends are separated in the bottoms.Degassing. The NMP solvent from thebottoms of the rectifier column is strippedfrom the heavy hydrocarbons, which areseparated as the distillate stream of thedegassing column. These heavy hydrocar-bons are then sent to a cooling column,where they are cooled by direct contactwith NMP solvent and fed to the rectifiercolumn through a gas compressor. In thedegassing column, the acetylenes areseparated as a side stream and washed ina scrubber before being purged. The NMPrecovered in the bottom stream is recycledto the extractive distillation columns.

Economic performanceAn economic evaluation of the processwas conducted, taking into considerationa unit processing a C4 stream to pro-duce 100,000 ton/yr of 1,3-butadieneerected on the U.S. Gulf Coast (the processequipment is represented in the simplifiedflowsheet below). The estimated total fixedinvestment for the construction of this plant isabout $50 million.

An important feature of the presentedtechnology is its low susceptibility to impuri-ties, rendering it adequate to process anyC4 stream, regardless of the butadienecontent. Also, plants using this technologyhave been shown to operate continuouslyfor more than four years.

In addition, the BASF NMP process tech-nology is currently applied in several plantsworldwide, as shown in Figure 2. n

Editor’s Note: The content for this column issupplied by Intratec Solutions LLC (Houston; www.intratec.us) and edited by Chemical Engineering.The analyses and models presented herein areprepared on the basis of publicly available andnon-confidential information. The information andanalysis are the opinions of Intratec and do notrepresent the point of view of any third parties.More information about the methodology forpreparing this type of analysis can be found,along with terms of use, at www.intratec.us/che.

1. First extractive column

2. Rectifier column

3. Second extractivecolumn

4. Propyne column

5. Butadiene distillationcolumn

6. Gas compressor

7. Cooling column

8. Degassing column

9. Scrubber

CW Cooling water

ST Steam

ST

CW

3

8

1

C4mixture

ST

2

Raffinate-1

CW

CW

CW

4

Propyne

CW

5

1,3-Butadiene product

Heavyends

ST ST

Acetylenes

7

6

9

CW

By Intratec Solutions

FIGURE 2.

BASF’s NMP pro-cess plants around

the world. Eachmark in the map

corresponds to an

existing facility

FIGURE 1. This 1,3-butadiene extraction process is similar to BASF’s NMP process



1,4-Butanediol from

Bio-Succinic Acid

The compound 1,4-bu-tanediol (BDO) is aversatile intermediate

for the chemical industry. Its

largest derivative product istetrahydrofuran (THF), whichis used to make spandexfibers, resins, solvents andprinting inks. The second larg-est product is polybutyleneterephthalate (PBT), whichis used to make high-perfor-mance materials, electronicsand automotive equipment.

BDO can be produced from different tech-nologies and raw materials. The conventionalmethod for manufacturing BDO is the Reppeprocess, starting from acetylene. Other pro-cesses use propylene oxide, maleic anhydride,1,3-butadiene or n-butane as starting points.The newest technologies being developed forBDO are bio-based pathways, which mostlyrely on bio-based succinic acid derived frombiomass or a sugar substrate.

Succinic acid is a platform chemical that canbe used to produce many products. However,its high cost of production from petroleum rawsources limits its use to specific applications(such as pharmaceuticals and food additives).According to the U.S. Dept. of Energy (DOE),bio-succinic acid is a renewable building-blockchemical with great potential for the future.

The processThe process depicted in Figure 1 was compiled

based on a U.S. patent published by BioAmberInc. (Montreal, Canada; www.bio-amber.com;U.S. patent no. 2011/0245515). The patentdiscloses details about the initial reaction,while the separation process was conceived byIntratec and is based on well-known practices.Hydrogenation reaction. A solution of bio-succinic acid in water is pre-heated and sent tothe fixed-bed hydrogenation reactor. Hydrogenis compressed and also fed in excess to thereactor. The process uses a bimetallic catalystconsisting of metals (including ruthenium, rhe-nium, tin and others) on a carbon support.

The exothermic reaction product is mostlyBDO (selectivities to BDO of more than 90%

are reported in literature). Side products includeTHF, gamma-butyrolactone (GBL), and linearalcohols (n-butanol, n-propanol). The amountof each product depends on the catalyst andoperational conditions used.

The bio-succinic acid is totally reacted,and the unreacted hydrogen is compressedand recycled back to the reactor. The productstream is cooled and sent to flash vessels torecover unreacted hydrogen and eliminatevolatile materials.Purification. The product from the flash vesselsis sent to a series of distillation columns toseparate BDO from side products. In the firstcolumn, BDO, GBL and water are removedfrom the bottoms and then sent to the dryingcolumn, where most of the water is removed.The column’s top fraction is sent to a lightscolumn, to recover side-products as fuel.

The recovered water is sent to a tank for fur-ther disposal or to be recycled to an integratedbio-succinic acid production unit (see Figure

2). The drying column’s bottoms are sent to aBDO recovery column and further to a heaviescolumn, where a stream of 99.5 wt.% BDO isfinally achieved.

Economic performanceAn economic evaluation of the process wasconducted based on data from the fourth quar-ter of 2013. The following assumptions weretaken into consideration.• A 55,000-ton/yr BDO production unit (the

process equipment is represented in Figure 1)• Storage capacity equal to 20 days of opera-

tion for BDO and no storage for bio-succinicacid

The estimated capital investment (includingtotal fixed investment, working capital andother capital expenses) to build a plant wouldbe $100 million in the U.S. Gulf Coast regionand $140 million in Brazil.

Process perspectiveIt may be an interesting consideration tocontrol the entire supply chain through theintegration of a bio-succinic-acid productionunit with a BDO unit (see Figure 2). Also, thebio-BDO produced must include a premiumprice over its petrochemical counterpart. Bothmeasures are important to make the ventureeconomically feasible.

In addition, it is also an important research-and-development goal to find hydrogenationcatalysts that have the following qualities:• Improved tolerance to impurities generated in

the fermentation process (that produces bio-succinic acid). This can help reduce the costsof bio-succinic acid purification

• Increased selectivity of bio-succinic acid toBDO, reducing the formation of side productsand the costs with BDO purification

Both achievements can reduce investment andproduction costs. n

Editor’s Note: The content for this column issupplied by Intratec Solutions LLC (Houston; www.intratec.us) and edited by Chemical Engineering.The analyses and models presented herein areprepared on the basis of publicly available andnon-confidential information. The information andanalysis are the opinions of Intratec and do notrepresent the point of view of any third parties.More information about the methodology forpreparing this type of analysis can be found, alongwith terms of use, at www.intratec.us/che.

Bio-succinic acid in water

ST

Hydrogen

Lights

as fuel

CW

CW

Water for

disposal or

recycle

Heavies

BDO

1

2

3

STCW 9 10

1) Hydrogenation reactor

2) High-pressure vessel

3) Low-pressure vessel

4) High-pressure column

5) Light-ends column

6) Drying column

7) BDO recovery column

8) Heavies column

9) Cooling tower

10) Boiler

CW Cooling water

ST Steam

CW

ST

4

CW

ST

6

CW

ST

7

CW

ST

8

CW

ST

5

Offgas

LightsMilling,

treatment and

sugar inversion

Sugarcane

FermentationCell separation,recovery and

purification

Hydrogenationreaction Purification BDO

Water

Bio-

succinicacid

CornMilling and

saccharification

Glucose

Hydrogen

Bio-succinic acid unit 1,4-Butanediol unitBiomass source unit

Heavies

or


FIGURE 2. It is possible to integrate 1,4-butanediol and bio-succinic-acid production units

IGURE 1. The 1,4-butanediol process from bio-succinic acid that is shown here

s based on patent disclosures from BioAmber and well-known practices



A solution-based

route to LLDPE

Polyethylene (PE) is the world’s largest-volume commodity polymer. Along withhigh-density (HDPE) and low-density

(LDPE) polyethylene, linear low-density poly-ethylene (LLDPE) is one of the three main typesof PE. The global market for LLDPE is over 20million metric tons per year, corresponding toabout 30% of the total PE produced.

In LLDPE production, three major types oflow-pressure technologies are used: slurry,solution and gas-phase processes.

The processLLDPE is produced by copolymerization ofethylene with alpha-olefins using Ziegler-Nattacatalysts. The most common co-monomers usedin LLDPE production are 1-butene, 1-hexeneand 1-octene.

Figure 1 illustrates the process for butene-based LLDPE production via a solution technol-ogy similar to Nova Chemicals’ (Calgary,Alta.; www.novachem.com) solution-phasetechnology, known as SclairTech (Figure 2).The process shown is a swing process, whichis also capable of producing different LLDPEand HDPE grades by utilizing other alpha-olefins as co-monomers.

The process can be divided into four mainoperation areas: purification and catalyst prep-aration; reaction; distillation; and finishing.Purification and catalyst preparation. The cyclo-hexane solvent, ethylene and comonomers aresent to fixed-bed adsorption systems to remove

water, oxygen and other polar impurities. Thecatalysts used in the process are based onmixtures of titanium and vanadium compounds,in conjunction with aluminum alkyls cocatalysts.These components are mixed with solvent andpumped to the polymerization reactor.Reaction. Ethylene and 1-butene comonomer(in case of butene-based LLDPE) are dissolvedin cyclohexane solvent and sent to the reactionstep. The polymerization is carried out in asolution phase, at a temperature above themelting point of the resulting polymer. The reac-tion system consists of a tubular reactor and acontinuous-stirred-tank reactor (CSTR). The lowresidence time of the reactors enables a high

flexibility for grade transitions, as well as ver-

satility for the production of resins with a widerange of densities, melt indexes and molecularweight distributions. The reactor output streamis fed into separator vessels, where unreactedethylene and co-monomers, solvent and anyother volatile matter are separated from the PE.The polymer is sent to the finishing section whilethe light stream moves to the distillation system.Distillation. The distillation step comprisesfive distillation columns in charge of recover-ing the unreacted ethylene and co-monomers;recovering the solvent; purging impurities,such as oligomers (also called grease), catalystand deactivators residues; and avoiding the

buildup of inert components, such as isomersof the co-monomer.Finishing. The resulting polymer from the reac-tion area is fed into an extruder, which is usedto incorporate the required additives, and topelletize the polymer. The product is then sentto the product blending and storage stage.

Economic performanceAn economic evaluation of the solution-phaseLLDPE process was conducted based on datafrom the fourth quarter of 2012. The followingassumptions were taken into consideration:• A 350,000 ton/yr unit erected on the U.S.

Gulf Coast (the process equipment is repre-

sented in the simplified flowsheet)

• Storage of products is equal to 20 daysof operation, and there is no storage forfeedstock

• Outside battery limits (OSBL) units con-sidered: steam boilers, cooling towers,propylene refrigeration system, heat-transferfluid unit, control room and administrativebuildings

The estimated capital investment (includingtotal fixed investment, working capital andother capital expenses) to build the LLDPE plantis about $220 million, and the operating costfor butene-based LLDPE production is about$1,220/ton.

The swing process depicted here allowsmanufacturers to participate in major PEmarket segments by producing both LLDPE andHDPE resins. Thus, the producers can select thebest product mix, aimed at premium marketswith higher margins. ■

Edited by Scott Jenkins

Editor’s Note: The content for this column is suppliedby Intratec Solutions LLC (Houston; www.intratec.us)and edited by Chemical Engineering. The analysesand models presented herein are prepared on the ba-sis of publicly available and non-confidential informa-tion. The information and analysis are the opinions ofIntratec and do not represent the point of view of anythird parties. More information about the methodol-ogy for preparing this type of analysis can be found,

along with terms of use, at www.intratec.us/che.

1. Adsorption system2. Tubular reactor3. CSTR polymerization

reactor4. Separators5. Extruder and pelletizing6. Low boiler column

7. Ethylene column8. Comonomer column9. Solvent column10. Grease column

11. Refrigeration unit12. Heat transfer fluid unit13. Cooling tower14. Boiler

HF Heat transfer fluid

CW Cooling water RF Refrigerat ion fluid

ST Steam BFW Boiler feed water

EthyleneButene

Cyclohexane

Catalyst and

cocatalyst

CW

HF

CW

HF

RF

ST

CW

ST

to FuelLLDPE

Additives ST

HF

CW

to Fuel

to Fuel

BFW

ST

BFW

ST

Deactivator

CW

RF

HF ST

CW11

9

76

5

4

3

2

1

8

10

12

13

14

FIGURE 1. Solution-phase LLDPE production process similar to Nova Chemicals’ SclairTech

Up to 149,000 ton/yr

From 150,000 to 249,000 ton/yr

From 250,000 to 349,000 ton/yr

At least 350,000 ton/yr

FIGURE 2.

Each mark on the mapcorresponds to anexisting SclairTechLLDPE plant. Thenominal capacityof each plant followsthe legend below




Bio-based adipic acid

Adipic acid is a dicarboxylic acid thatis widely applied to the productionof nylon, polyurethanes, plasticizers

and other polymers. Its main use is as oneof the raw materials for the manufacture ofthe polyamide nylon 6,6. In 2012, 80% ofthe total adipic-acid demand was directedto this application.

The traditional production method foradipic acid is based on the oxidation ofcyclohexane, a petroleum-based feedstock.However, adipic acid is among the growinggroup of industrial compounds for whichmore-sustainable production methods havebeen developed. A bio-based route toadipic acid has emerged as an alternativeto petroleum-derived adipic acid. Bio-adipicacid can be produced from renewable feed-stocks, such as sugar and plant-based oils.

The processAdipic acid is produced by a microorgan-ism-based fermentation process that uses asucrose solution as the feedstock (Figure 1).The process described below was compiledbased on information available in the pub-lished literature.Sugar inversion. Sucrose is hydrolyzed tofructose and glucose, which are the com-pounds actually consumed in fermentation.The reaction occurs in an acidic medium,obtained by adding HCl. After inversion,the pH is adjusted with Na2CO3, and theresulting stream is cooled before being sentto the fermentation stage.Fermentation. The sugars are converted toadipic acid through an aerobic fermentation.The microorganism culture is prepared forinoculation in a two-stage seed train. An am-monia (NH3) solution from the recovery areais added to the fermenters to maintain a pHof about 7. The adipic acid produced inthe fermentation step reacts with ammoniumhydroxide, forming diammonium adipate.

Cell separation. The fermentation broth issent to agitation tanks and centrifuges inorder to completely remove cell biomass.

Ultrafiltration is used to remove remain-ing contaminants, such as cell debris andprecipitated proteins. The aqueous solution,now free of microorganisms, is then sent tothe product-recovery area.Product recovery. The solution is sent to theconcentration column, where diammoniumadipate is split into free ammonia andadipic acid. The overhead stream, whichcontains mostly ammonia and water, is sentto the ammonia-recovery column and, after-wards, is recycled to the fermentation step.The bottoms product stream contains dis-solved adipic acid that is crystallized at lowtemperatures. The slurry from the crystallizeris centrifuged, and part of the mother liquoris recycled to the concentration column. Thesolid adipic acid product is dried and sentto storage.

Economic performanceAn economic evaluation of the bio-basedadipic acid production process was con-ducted. The following assumptions weretaken into consideration:• A 83,000-ton/yr capacity plant erected

on the U.S. Gulf Coast• Storage of product is equal to 30 days

of operation, and there is no storage forfeedstock

• Outside battery limits (OSBL) units con-sidered: ammonia refrigeration system,cooling towers and steam boilers

The estimated capital investment (includingtotal fixed investment, working capital andother capital expenses) for the constructionof such a plant amounts to about $340million, and the operating cost is about$2,150 per ton of product, as depicted inFigure 2.

Due to the availability of low-cost feed-

stock, the bio-based process can be a goodalternative when compared to petroleum-based process for adipic acid production.Research activity to achieve higher product

yield and selectivity may eventually leadto the bio-based process becoming morecost-effective than the conventional petroleum-based process.

Aside from the possible economic advan-tages, the production of bio-based adipicacid can have an impact on environmentalissues. The bio-based process is more envi-ronmentally friendly and leads to an overall

reduction in oxides of nitrogen (NOx) andother greenhouse-gas emissions comparedto what is normally produced during thepetroleum-based adipic acid process. n

Editor’s Note: The content for this column is suppliedby Intratec Solutions LLC (Houston; www.intratec.us)and edited by Chemical Engineering. The analysesand models presented herein are prepared on the ba-sis of publicly available and non-confidential informa-tion. The information and analysis are the opinions ofIntratec and do not represent the point of view of anythird parties. More information about the methodol-ogy for preparing this type of analysis can be found,along with terms of use, at www.intratec.us/che.

50%

22%

28%

Fixed costs

Main utilities consumption

Net raw materials cost


FIGURE 2. Raw materials costs account for thehalf of the operating expenses for a bio-based

adipic acid unit

1. Hydrolysis reactor

2. Neutralization vessel

3. Air compressors

4. Seed trains

5. Fermenter

6. Cells centrifuge

7. Concentrationcolumn

8. Ammonia recoverycolumn

9. Crystallizer

10. Dryer

11. NH3 refrigerationunit

12. Cooling tower

13. Steam boiler

RF Refrigeration fluid

CW Cooling water

ST Steam

Micro-organismsseed

Exhaustgases

SucroseSolution

HCl

Na2CO3

Air

CW

CW

To wastetreatment

CW

Water make-up NH3 make-up

Adipicacid

To wastetreatment

RF

ST

ST

CW

CW

To wastetreatment

RF CW11

10

9

7

651

2

3

4 8

12

13ST

FIGURE 1. Bio-based adipic acid production is based on fermentation, ultrafiltration and crystallization processes



Production of Bio-based

Succinic Acid

Succinic acid is an important industrialchemical that is traditionally derivedfrom petroleum feedstocks. It can be

used in several areas — from high-valueniche applications, such as food, bever-ages and pharmaceuticals, to large-volumeapplications, such as plasticizers, resins andcoatings. However, the current market forsuccinic acid is limited to specialized areasbecause of its costly production process.

Bio-based succinic acid, produced fromrenewable feedstocks, appears to be an al-ternative to the petroleum-derived chemicaland may have advantages surrounding con-cerns about rising atmospheric greenhousegas (GHG) concentrations and volatilecrude oil prices. Bio-succinic acid is chemi-cally equivalent to conventionally producedsuccinic acid, and is suitable for the sameapplications. For more on bio-based suc-cinic acid, see Chem. Eng., August 2013,pp. 14–17.

The processIn the process depicted in Figure 1, succinicacid is obtained via microbial fermenta-tion of glucose from sugarcane juice. Theprocess steps and equipment were compiledbased on publically available informationfrom the scientific literature.Sugar inversion. The sugar consumed in thisfermentation-based process is glucose. It isobtained by the inversion of sucrose throughits hydrolysis to glucose and fructose. This isachieved by subjecting the juice containingsucrose to an acidic medium by additionof hydrochloric acid (HCl). After inversion,the pH is adjusted with sodium carbonate(Na2CO3) and the stream is cooled beforebeing sent to the fermentation step.Fermentation. Glucose is converted tosuccinic acid through an anaerobic fer-mentation supplied with carbon dioxide.The microorganism culture is preparedfor inoculation in a two-stage seed train.Carbon dioxide is compressed, filtered and

cooled before being fed into the fermenter.A solution of ammonia (NH3) is added tothe fermenters to maintain a pH level of

about 7.Cell separation. The fermentation broth issent to surge tanks and then to centrifugesin order to completely remove cell biomass.Ultrafiltration is used to remove the remainingcontaminants, such as cell debris and precipi-tated proteins. Following that, the aqueoussolution, free of microorganisms, is sent tothe product recovery and purification area.Recovery and purification. In this stage, adecolorization process is performed withthe addition of activated carbon. Then, themixture is filtered to separate the activatedcarbon. The aqueous phase is sent to avacuum-distillation column, where volatilebyproducts and some water are removed.This distillation leads to a more concen-trated succinic acid stream. The succinicacid product is then crystallized at lowtemperature. HCl is used to adjust pH,allowing selective crystallization of succinicacid. Succinic acid crystals are washed withwater in a filter and dried with hot air. Partof the succinic acid product that is lost in thefiltrate is recycled to the distillation column.

Economic performanceAn economic evaluation of the bio-basedsuccinic acid process was conducted. Thefollowing assumptions were taken intoconsideration:• A 77,000 ton/yr unit erected on the U.S.

Gulf Coast (the process equipment is rep-resented in the simplified flowsheet below)

• Storage of product is equal to 30 daysof operation, and there is no storage forfeedstock

• Outside battery limits (OSBL) units consid-ered: steam boilers, cooling towers andammonium refrigeration system

The estimated capital investment (includingtotal fixed investment, working capital andother capital expenses) is about $260 mil-

lion, and the operating expenses are about$2,200/ton of product (Figure 2).

Beginning with high-value applica-tions, bio-succinic acid has the potential tobecome a key building block for commod-ity chemicals, since consumers of end-useproducts are increasingly demanding moreeco-friendly products. Also, further develop-ment in bio-based succinic acid processeswill likely lead to more cost-effective produc-tion, resulting in an increase in the demandfor this product compared to the costlypetroleum-based succinic acid.

Based on these features, several newprojects for producing bio-based succinicacid have begun operations, and othersare planned in the near future, thus grow-ing the worldwide production capacity forthis key industry chemical. n

Editor’s Note: The content for this column is supplied byIntratec Solutions LLC (Houston; www.intratec.us) andedited by Chemical Engineering. The analyses andmodels presented herein are prepared on the basisof publicly available and non-confidential information.The information and analysis are the opinions of In-tratec and do not represent the point of view of anythird parties. More information about the methodologyfor preparing this type of analysis can be found, alongwith terms of use, at www.intratec.us/che.

CW

Microorganismsseed

CW

ST

Canejuice

HCl

Na2CO

3

CO2

Activatedcarbon

To wastetreatment

Lights

Water

HClSuccinicacid

Exhaustgases

RF

NH3

To wastetreatment

CW

Hot air

7

CW

Processwater

Purge

RF

ST

CW

1. Hydrolysis reactor


3. CO2 compressors

4. Seed train

5. Fermenters

6. Cells centrifuge

7. Vacuum column

8. Crystallizer

9. Dryer

10. NH3 refrigeration unit

11. Cooling tower

12. Steam boiler

RF Refrigeration fluid CW Cooling water

ST Steam12

1110

8

76

9

54

3

2

1

FIGURE 1. Bio-succinic acid is produced via the fermentation of glucose

0

500

1,000

1,500

2,000

2,500

$ / t o n o

f p r o d u c t

Net raw materials costs

Main utilities consumptions

Fixed costs


FIGURE 2. The operating expenses for bio-succinic acid are broken down here



Producing Ethanol from

Lignocellulosic Biomass

Pressing global issues, such as climatechange and declining energy security,have led to an increasing interest in

the use of renewable feedstocks for fuel pro-duction. These feedstocks include lignocellu-losic materials, such as waste biomass fromthe agriculture and forest industries.

Ethanol is commonly produced byfermentation of sugars from agriculturalfeedstock — mostly corn and sugarcane.Several processes to produce ethanol byfermenting lignocellulosic materials havebeen developed. Two such processes havebeen introduced by American Process Inc.(Atlanta, Ga.; www.americanprocess.com),one of which was covered in this columnlast year (Chem. Eng., November 2013, p.26). The other cellulosic ethanol process isdiscussed below.

The processThe process for cellulosic ethanol productiondepicted in Figure 1 is similar to the AVAPprocess described in the World IntellectualProperty Organization (Geneva, Swit-zerland; www.wipo.int) patent publishedby American Process Inc. (W.O. Patent2011/044378A1).Pretreatment. The biomass is heated witha solution of SO2, ethanol and water andthen fractionated into its three lignocellulosiccomponents: hemicellulose, cellulose andlignin. Hemicellulose and lignin are partiallydissolved, while cellulose remains insoluble.Cellulose is washed and the filtrate is sent tothe stripping column, where SO2 and etha-nol are recovered in the overhead stream,which is condensed in an evaporator andsent to a make-up vessel. The column bot-toms material is concentrated in the evapo-rator, generating vapor that is compressedand used as steam supply for the evapora-tor and for the column’s reboiler.Cellulose hydrolysis and fermentation. Cel-lulose is hydrolyzed into glucose monomers

by an enzymatic reaction.Insoluble lignin is separatedand conducted to the hemi-

cellulose hydrolysis areaand glucose is fermented toethanol. The broth is cen-trifuged, then the cells aretreated with sulfuric acid andrecycled to the fermenters.Hemicellulose hydrolysisand fermentation. Theconcentrated stream from thepretreatment area is mixedwith the lignin from the enzymatic hydrolysisand pumped to the auto-hydrolysis reactor,where hemicellulose is hydrolyzed to mono-mer sugars, such as xylose. The hydrolyzateis neutralized with lime to precipitate thesoluble lignin. Insoluble lignin is separatedand burned to generate steam and energyto supply plant demands. The monomersugars are fermented to ethanol.Distillation. The broths from the fermenta-tion steps are mixed and conducted to theconcentration column, where a vapor sidestream with 50% ethanol is obtained. Thebottoms material from this column is used towash cellulose. The side stream is fed to therectifying column, generating an overheadstream containing about 93% ethanol. Theethanol product stream undergoes dehydra-tion in a molecular sieve unit, leading toethanol product with 99.5% purity.

Key research featuresRegarding its process, American Processsays the following:• The application of enzymes only for cel-

lulose hydrolysis requires a lower doseof enzymes when compared to processesconverting both cellulose and hemicellu-lose by enzymatic hydrolysis

• The process offers an energy-efcientrecovery of cooking chemicals, resulting inlow expenses with chemical make-up

Economic performanceAn economic evaluation of the process wasconducted, taking the following assumptionsinto consideration:• A biomass processing unit producing 58

million gallons of anhydrous ethanol per year, built on the U.S. Gulf Coast

• Storage of products is equal to 30 daysof operation, and there is no storage forfeedstock

The estimated capital expenses (includingtotal fixed investment, working capital andother expenses) for the construction of thisplant are about $700 million.

An interesting application to the pre-sented technology is its integration withthe pulp-and-paper industry. In this case,the process can fractionate wood into itsprimary components (cellulose, hemicel-

lulose and lignin). Then, each fraction canbe used in different applications, targetingeach toward the most profitable market(Figure 2). n


1. Cooking digester

2. Cellulose washers

3. Stripping column

4. Make-up vessel

5. Two-effect evaporator

6. Enzymatic hydrolysis reactor

7. Cellulose fermenters

8. Auto-hydrolysis reactor


10. Hemicelluloses fermenters

11. Concentration column

12. Rectifying column

13. Molecular sieve unit

14. Cooling tower

15. Energy cogeneration unit

CW Cooling water

ST Steam

SO2

H2O

Ethanol

Biomass

CW

Ventgases

Sulfuric acid

Lime Ventgases

Lignin

CW

ST

EthanolST

CW

To wastetreatment

Ventgases

ST

Electricity

Enzyme

CW

To wastetreatment To wastetreatment

CW

13

14

12

11

10

15

9

8

67

5

3

1

2

4

Recovered chemicals condensate

Cellulose Hemicellulose Lignin

Pulp and paperproducts

Ethanol Energy

AVAP

process (biomass fractionation)

Wood


FIGURE 2.

Wood fraction-ation for the

manufacture of

different finalproducts

FIGURE 1. Cellulosic ethanol production process similar to the AVAP process from American Process Inc.



Ethanol from the Direct

Gasification of Biomass

Increasing global demand for energy anda drive to develop renewable energysources have focused research and

development attention on biomass (includingwood, agricultural residues and municipalwastes) as a source of energy and products.Biomass can be converted into bioenergy viabiochemical and thermochemical means. Inbiochemical routes, the biomass is convertedinto liquid or gaseous fuels by fermentationor anaerobic digestion. In thermochemicalconversion, biomass is converted into gasesvia a number of methods, and the gasesare then either used directly or synthesizedinto the desired chemicals. Thermochemicaltechnologies include combustion, gasifica-tion and pyrolysis. This column describes agasification process to produce ethanol.

The processIn the process of producing ethanol via directgasification (Figure 1), biomass is con-verted into synthesis gas (syngas) by partialoxidation and the resulting stream of gasesis partially burned to produce heat for theprocess and partially reacted to produce liq-uid alcohols. The process can be divided intothe following areas: feedstock handling andgasification, gas cleanup and conditioning,alcohol synthesis and alcohol separation.Feedstock handling and gasification. Thebiomass is dried with fluegas from the fuelcombustor in the biomass drier. The driedbiomass is fed into the oxygen-blown directgasifier with steam and high-pressure oxy-gen from an air-separation unit.

In the presence of steam and oxygen athigh temperature, biomass decomposes intosyngas — a mixture of mostly H2 and CO,with CH4, CO2, light hydrocarbons andwater — and tar, ash and char. The com-bustion of a portion of the biomass suppliesthe heat for the endothermic reactions. Acyclone is used at the exit of the gasifier toremove residual solids from the syngas.

Gas cleanup and conditioning. Syngasfrom the gasifier is sent to the fluidized-bedtar reformer, where the hydrocarbons are

converted to CO and H2. The deactivatedreforming catalyst is separated from theeffluent syngas in a cyclone, regeneratedand sent back to the tar reformer. The hotsyngas is used for steam generation in heatexchangers and is then sent to the quenchtower. The cooled gas from the quench issent to the amine scrubber unit for the re-moval of CO2 and H2S. The CO2 is ventedto the atmosphere and H2S is reduced tosulfur before disposal.

Alcohol synthesis. The treated gas is com-pressed and sent to the alcohol synthesisreactor, a fixed-bed system using a MoS2-based catalyst. Carbon monoxide and H

2

combine through a series of reactions toform a mixture of alcohols, with ethanolbeing the main product. The reactionproducts are cooled so that the mixture ofalcohols is condensed and separated fromthe unconverted syngas. The unconvertedsyngas is recycled to the tar reformer witha small purge.

Alcohol separation. The condensed streamfrom the alcohol synthesis reactor is vapor-ized and dehydrated by molecular sieves.The dehydrated alcohol mixture is fed tothe higher alcohols column for separationof higher alcohols as a byproduct. Theoverhead from the column is then sent to themethanol column, where ethanol product isobtained. The column overhead, contain-ing mainly methanol is recycled back to thealcohol synthesis reactor.

Economic performanceThe capital expense of the process wasestimated based on data from the secondquarter of 2013 (Figure 2). The followingassumptions were taken into consideration:• A biomass processing unit producing 30

million gallons of ethanol per year erectedon the U.S. Gulf Coast

• The process equipment is represented inthe simplified flowsheet in Figure 1

• In the process, steam is generated throughheat recovery from hot process streams.The power requirements for the process isgenerated in turbines driven by the steamproduced

• Storage capacity equal to 30 days ofoperation for ethanol product

Ethanol production from biomass gasifica-tion is a promising technology that canuse a wide range of low-cost feedstocks toproduce liquid fuels. The main challenges

for the technology are the development ofmixed-alcohol catalysts, reforming and thehigh investment associated with the construc-tion of a commercial plant. n


1. Biomass drier

2. Air separation unit3. Gasifier

4. Fuel combustor

5. Tar reformer

6. Catalyst regenerator

7. Quench column

8. Amine absorber

9. Amine regenerator

10. H2S removal system

11. Compression unit

12. Alcohol synthesis reactor

13. Molecular sieve unit

14. Higher alcohols column

15. Methanol column

16. Power generation

17. Cooling towerBFW Boiler feed water

CW Cooling water

ST Steam

Air

Biomass

Flue gas

ST Solidwaste

Air

O2

Catalyst ST

BFW

ST

To wastetreatment

CW

CW

CWCO2

Sulfur

CW

ST

Flue gas

Catalyst purge

STST

CWCW

Higheralcohols

by-productEthanol

ST

Electricity

.

CW

ST

ST

BFW

15

17

16

14

13

12

10

11

98

3

2

1

4

56

7

FIGURE 1. Ethanol production from direct gasification of biomass

$ , m i l l i o n s

300

250

200150

100

50

0

Total fixed investmentOther capital expensesWorking capital


FIGURE 2. Capital expenditures to build a plantfor ethanol



Ethylene Production

via Ethanol Dehydration

Ethylene is arguably the most importantbuilding block of the petrochemicalindustry, frequently produced via steam

cracking of a range of petroleum-based feed-stocks. Rising oil prices and concerns aboutglobal warming have motivated research intoethylene manufacture from renewable sources.

The “green” ethylene made from ethanol(for example, from corn, sugarcane or lignocel-lulosic biomass) represents a chemically identi-cal alternative to its petrochemical equivalent,and can contribute to the overall reduction ofgreenhouse gas emissions.

Ethanol dehydrationThe polymer-grade (PG) ethylene production viaethanol dehydration process depicted in the fig-ure below was compiled based on a U.S. pat-

ent published by BP Chemicals (London, U.K.;www.bp.com; U.S. patent no. 8,426,664).Since 2008, BP has established more than tenpatents on this subject.

Ethanol dehydration has an overall endother-mic equilibrium. The ethylene yield is favoredby higher temperatures, while lower tempera-tures favor production of diethyl ether (DEE).The major process steps are outlined below.Treatment. Fresh ethanol is combined with therecycled ethanol and DEE from the ethylenecolumn and sent to the aldehyde removal col-umn, which removes the aldehydes that camefrom the fresh ethanol feed, as well as thealdehydes and C4 hydrocarbons generated in

the reaction step.Reaction. The bottom stream from the alde-hyde removal column is sent to the feed vapor-izer and superheater before being sent to thereactors. The process uses fixed-bed reactorswith a heteropolyacid catalyst supported bysilica. The product stream from reaction iscooled and sent to the purification section.Purification. The reactor’s outlet streamprimarily consists of unreacted ethanol, water,ethylene and DEE. The stream is sent to afirst distillation column to separate mainly theethylene from the water. The ethylene-richoverhead stream is then compressed, driedand sent to an ethylene purification column.

The bottom stream from the ethylene-waterseparation column is then sent to a dewateringcolumn, to separate water. Streams containing

mainly ethanol and DEE are recycled back tothe treatment section.

Key research basisThe key information used as input to developthe synthesis of the process is presented below.The final process design was based on reason-able values from the ranges presented here.Since the technology is in its embryonic stages,and no commercial plant has been built usingthis concept thus far, the design depicted hereis subject to changes.

Reactor temperature 200 to 250°CReactor pressure 10 to 28 bars abs

Ethanol conversion to ethylene(per pass) 20 to 60%Feed ethanol water content 1 to 5 wt.%Diethyl ether in reaction feed 50 to 85 wt.%Catalyst-bed temperature profile <10°C

• According to BP, the formation and recyclingof DEE thermodynamically favor the separa-tion of water, and result in higher ethyleneselectivity

• Moderate reaction conversion and mildtemperature and pressure conditions areused in order to improve the selectivity forthe desired products.

Economic performanceAn economic evaluation of the process wasconducted based on data from the secondquarter of 2012 with the U.S. Gulf Coast asthe location. The following assumptions wereassumed for the analysis:• A 190,000-ton/yr chemical production unit

(all equipment represented in the simplifiedflowsheet below)

• Storage capacity equal to 20 days of opera-tion for ethanol and no storage for ethylene

• Outside battery limits (OSBL) units con-sidered: steam boilers, cooling towers,propylene refrigeration system, controlroom and administrative buildings

The estimated capital investment (includingtotal fixed investment, working capital andother capital expenses) to build an ethanol

dehydration unit would be between $200and $280 million, and the operating costwould be about $1,650/ton. Given thatthe petroleum-based ethylene sales price islower than the green (via ethanol dehydra-tion) ethylene cost of production, the greenethylene sales price must include a premiumof more than 50% over petrochemically de-rived ethylene to make the venture economi-cally feasible.

Research perspectiveStrategic, financial and technical insightsabout this research into the ethanol dehydra-tion process inlcude the following:

• Marketing and sales teams must nd niche-market consumers willing to pay more for anenvironmentally friendly product (for example,plastic containers for the cosmetics industry)

• Since ethanol raw material prices are dictat-ed by markets such as food and fuel, it maybe an interesting consideration to control theentire supply chain through the integration ofthe ethanol dehydration unit with an ethanolproduction unit

• The published patents describe problemsrelated to the temperature profile of thecatalyst bed. Researchers must focus on de-veloping a catalyst with improved toleranceto the temperature decrease inherent in the

endothermic nature of the dehydration reac-tion. Higher conversion and higher selectivityfor ethylene would minimize recycling loopsand, consequently would reduce both capitaland operating costs ■


Editor’s Note: The content for this column is suppliedby Intratec Solutions LLC (Houston; www.intratec.us)and edited by Chemical Engineering. The analysesand models presented herein are prepared on the ba-sis of publicly available and non-confidential informa-tion. The information and analysis are the opinions ofIntratec and do not represent the point of view of anythird parties. More information about the methodol -ogy for preparing this type of analysis can be found,along with terms of use, at www.intratec.us/che.

ST

CWEthanol

ST

CW

ST

RF

ST

CW

PGethylene

Water

FU

1

2 5

6

7

8

9 10RF

1) Aldehyde removal column

2) Feed vaporizer

3) Feed superheater

4) Dehydration reactors

5) Cooler

6) Ethylene-water column

7) Compressor and drying

8) Ethylene column

9) Dewatering column

10) Propylene refrigeration unit

FU FuelCW Cooling water

RF Refrigerat ion fluid

ST Steam

Diethyl ether recycle

Ethanol, water and diethyl ether recycle

3

4

By-productsto fuel




Hydrogen Cyanide

Production

H ydrogen cyanide (HCN) is a chemicalprecursor used in the production ofseveral industrially relevant compounds,

such as adiponitrile for nylon production,methyl methacrylate for polymer manufactur-ing, sodium cyanide for gold recovery and forthe production of methionine, which is used asa feed additive.

HCN is mainly produced though the Andrus-sow process, named for developer LeonidAndrussow. The method involves reactingammonia, natural gas and air over a platinumcatalyst to form HCN. Alternatively, HCN canbe produced by the BMA process, which usesammonia and natural gas only.

The processHCN production via the Andrussow processis depicted in Figure 1 and described below,

based on information available in the litera-ture. The process can be divided into threemain areas: reaction, ammonia recovery andproduct purification.Reaction. Natural gas, ammonia and air arefed into the reactor, where HCN is formedthrough a catalytic reaction. The productstream, containing HCN and unreacted am-monia, must be cooled down to avoid HCNdecomposition. This is accomplished in awaste-heat boiler located below the reactor.The waste-heat boiler generates steam that canbe used elsewhere in the process. Ammonia recovery. The product stream is fedto the bottom of the ammonia absorber, wherephosphate is used to absorb ammonia. The

overhead stream of this column, which containsmainly HCN, is sent to the purification section.The bottoms stream, containing ammonia anda small amount of HCN, is heated and sent tothe first HCN stripper.

In this column, HCN is separated in the over-heads and recycled to the ammonia absorber.The bottom stream is fed to the top of the am-monia stripper, where phosphate is separatedin the bottoms. Part of this stream is mixed withphosphoric acid and recycled to the top of theammonia absorber. Ammonia, now free ofphosphate, is separated as the overhead streamof the ammonia stripper. The ammonia is thendirected to the dryer, where it is concentrated

and recycled to the reaction section.Product purification. The HCN-containingstream is fed to the HCN absorber, whereHCN is absorbed in cold water. The overheadvent gas is incinerated, since it can still containsome HCN. The bottom stream from the HCNabsorber is directed to the second HCN strip-per, where the product is separated from wa-ter, and then purified in the HCN rectificationcolumn to obtain 99.5%-pure HCN product.

Economic performanceAn economic evaluation of the process wasconducted, taking the following assumptionsinto consideration:• A 22,800 ton/yr unit erected on the U.S.Gulf Coast (the process equipment is repre-sented in the simplified flowsheet below)• Outside battery limits units considered arethe cooling tower and the refrigeration system• No storage was considered

The capital investment (including total fixedinvestment, working capital and other capitalexpenses) for the construction of this plant isestimated to be about $80 million, while theoperating expenses are estimated to be about$1,070 per ton of product. The variation of the

operating expenses according to plant capac-ity is depicted in Figure 2.

Global perspectiveHCN is a toxic compound and is danger-ous to the environment as well as to humanhealth. Additionally, it is flammable and, if notproperly stored, may undergo a polymeriza-tion reaction that can be explosive. For thesereasons, HCN storage and transportation mustbe avoided, so HCN is essentially a non-tradable product. Companies that use HCN asa feedstock in their processes must produce itthemselves or have it supplied to them by anintegrated HCN facility. Also, companies maypurchase HCN from nearby acrylonitrile plantsthat produce it as a byproduct. n

Editor’s Note: The content for this column issupplied by Intratec Solutions LLC (Houston; www.intratec.us) and edited by Chemical Engineering.The analyses and models presented herein areprepared on the basis of publicly available andnon-confidential information. The information andanalysis are the opinions of Intratec and do notrepresent the point of view of any third parties.More information about the methodology forpreparing this type of analysis can be found, alongwith terms of use, at www.intratec.us/che.

1) Reactor waste-heat boiler

2 ) Ammonia absorber

3) First HCN stripper

4 ) Ammonia stripper

5 ) Ammonia dryer

6 ) HCN absorber

7 ) Second HCN stripper

8 ) HCN rectification column

9 ) Refrigeration unit

10) Cooling tower

BFW Boiler feed water

CW Cooling water

ST Steam

RF Refrigeration fluid

3

4

5

78

6

BFW

ST

CW

1

2

Natural

Gas

Air

NH3

make-upST

Phosphoric acid

To wastetreatment

To wastetreatment

To wastetreatment

CW

To waste

treatmentTo waste

treatment

ST

RF

ST

CW

HCN

ST ST

9

10CW

Vent gasto incineration

ST

RF

ST


FIGURE 2. Operating expenses for HCN production variation with plant capacity

FIGURE 1. The HCN production scheme shown here is based on the Andrussow process

1250

1200

1150

1100

1050

1000

950

900

O p e r a t i n g

e x p e n s e s

( $ / t o n

)

Plant capacity (thousands ton/yr)

10 15 20 25 30 35 40 45 50



Methanol-to-Olefins Process

Ethylene and propylene are the mostimportant intermediates in the petro-chemical industry. Globally, they are

produced mainly by steam cracking of hydro-carbons, such as naphtha, propane and eth-ane. The methanol-to-olefins (MTO) process isan alternative approach to producing theselight olefins from methanol feedstock, whichcan be derived from other raw materials,including natural gas, coal or biomass.

The processFigure 1 illustrates an MTO process similarto one developed jointly by UOP LLC (DesPlaines, Ill.; www.uop.com) and NorskHydro A/S (Oslo, Norway; www.hydro.com). This process synthesizes olefins frommethanol using a SAPO-34-type zeolite cata-

lyst in a fluidized-bed reactor. The processcan be divided into three main areas: reac-tion and regeneration; quench, compressionand caustic wash; and product fractionation.Reaction and regeneration. Methanol feed isvaporized, mixed with recovered methanol,superheated and sent to the fluidized-bedreactor. In the reactor, methanol is first con-verted to a dimethylether (DME) intermediateand then converted to olefins with a veryhigh selectivity for ethylene and propylene.

During the reaction, coke accumulateson the catalyst, which is circulated to thefluidized-bed regenerator system. In the

regenerator, the coke is removed by combus-tion with air to maintain the catalyst activity.After leaving the reactor, the reacted streamexchanges heat with the reaction feed, inorder to recover the heat generated by theexothermic reaction.Quench, compression and caustic wash. The output from the reaction and regen-eration is quenched, where most of waterand unreacted methanol is removed.The water-methanol stream is sent to themethanol-recovery column, where methanolis recovered and recycled to the reaction.

The vapor stream from thequench stage is compressedand sent to a caustic wash

column for CO2 removal.The gas stream, free ofCO2, is then dried and sentto the product fractionation.Product fractionation. The dried stream from causticwash is sent to the de-etha-nizer column, where ethyleneis separated from C3 andlarger (C3+) hydrocarbonsas the overhead product.The ethylene-rich stream iscompressed and fed to theacetylene reactor for selective acetylenehydrogenation to ethane. The acetylenereactor effluent is sent to the de-methanizercolumn, where an ethylene-ethane mixture isseparated from a methane-rich stream. Theethylene-ethane mixture is then routed to theC2 splitter to obtain polymer-grade ethyleneproduct as the overhead and an ethane-richstream at the bottom.

The C3+ stream from the bottom of thede-ethanizer column is fed to the de-pro-panizer column, for C4+ separation as thebottom product. The de-propanizer columndistillate, mainly containing a propylene-propane mixture, is fed to the C3 splittercolumn. In the C3 splitter, polymer-gradepropylene is obtained as the overhead prod-uct, while the bottom stream is propane-rich.

Economic performanceThe total fixed investment of the processwas estimated based on data from thesecond quarter of 2013 for units erected onthe U.S. Gulf Coast and in China (Figure2). The following assumptions were takeninto consideration:• The plant capacity is 600,000 ton/yr of

light olefins (ethylene and propylene)

• The process equipment is represented in

the simplified flowsheet• The plant is constructed inside a chemicalcomplex. The methanol feedstock is locallyprovided and propylene and ethyleneproducts are consumed by nearby poly-olefins units. Therefore, no storage forproduct or raw material is required

The MTO process is an alternative ap-proach to producing ethylene and propyl-ene for chemical manufacturers that haveaccess to cheap sources of methanol, whichcan be supplied from catalytic conversionof synthesis gas. China’s low manufacturingcosts, as well as low investment require-ments, explains the recent plans to constructMTO plants in the country to monetizeexcess coal resources. In the U.S. on theother hand, methanol can be produced fromlow-cost shale gas. n


1. Reactor

2. Regenerator

3. Quench tower

4. Compression

5. Methanol recovery column

6. Caustic wash column

7. Deethanizer column

8. Acetylene reactor

9. Demethanizer

10. C2splitter

11. Depropanizer

12. C3splitter

13. Refrigeration unit

14. Cooling tower

15. Boiler

CW Cooling waterRF Refrigeration fluidST Steam

1

Methanol

Propane

RF

STCW

ST

Ethylene

PG PropyleneCW

STC

4+ byproduct

Tail gas

CW

STCW

STWater

Regen gas

Air

Causticsoda

Processwater

To waste

RF

ST

RF

ST

Ethane

12

11

13 14 15

109

RF STCW

8

7

74

2

3

1

FIGURE 1. Methanol-to-olefins technology similar to UOP/Hydro MTO process

400

350

300

250

200

150

100

50

0

U.S. Gulf Coast China

$ , m i l l i o n

Total fixed investment


FIGURE 2. Total fixed investment to build an MTO plant indifferent regions



Methanol-to-Propylene

Technology

Propylene has been established as animportant part of the global olefins busi-ness, trailing only ethylene in terms of

volume of product. Because inexpensive natu-ral gas from shale in the U.S. is increasinglyused as a feedstock for producing ethylene,lower quantities of three- and four-carbon ole-fins are available. For this reason, new chemi-cal processes for on-purpose propylene usinglow-cost raw materials are gaining importance.An example of such a process is methanol-to-propylene (MTP) technology.

The main raw material used in the MTPprocess is methanol that is produced fromsynthesis gas which, in turn, can be obtainedin large-scale from natural gas or coal.

The processThe MTP process consists basically of two

reaction steps: an initial one to dehydratemethanol to dimethyl ether (DME) on analuminum oxide catalyst, and a second oneto transform DME and methanol into a varietyof olefins, ranging from ethylene to octenes.However, using a zeolite-based catalyst (ZSM-5), the process yields mainly propylene. A setof purification columns is necessary to obtainthe polymer-grade (PG) propylene.

Figure 1 illustrates a process similar to theones licensed by Lurgi GmbH (the MTP pro-cess; Frankfurt am Main, Germany; www.lurgi.com) and JGC Corp. (Yokohama, Japan; www.jgc.co.jp) and Mitsubishi Chemical Corp.’s(Tokyo; www.mitsubishichemical.com) DTP

process, which can be divided into three mainareas: reaction and regeneration; quench andcompression; and product fractionation.Reaction and regeneration . In this stage,the methanol feed is vaporized, mixed withrecovered methanol and dimethyl ether (DME),superheated and sent to the DME reactor,where dehydration occurs.

The product is mixed with recycled hydro-carbons and steam before being fed into theMTP reactors. The reactors were designed withseveral stages to better approach isothermalconditions. Propylene synthesis is conductedin multiple reactors: while one set of reactorsconducts the reaction, the remaining ones arein regeneration or on stand-by mode. The recy-

cling of olefins increases propylene yield andabsorbs the heat generated in the reaction.

After leaving the reactors, the mixture iscooled and sent to the quench and compres-sion steps.Quench and compression . The output fromthe reactors is quenched, where most of thewater is removed. A portion of this water issent to the methanol-recovery column, whilethe remaining water is vaporized and used asdilution steam in the MTP reactors. The vaporstream from the quench stage is compressed,partially condensed and then separated intoliquid and vapor streams.Product fractionation . The liquid stream fromthe quench and compression stages is sent tothe debutanizer, to separate four-carbon C4hydrocarbons (and those with less than fourcarbons) and DME from those with greater

than five carbons (C5+ components). The C5+stream from the column bottom is then sepa-rated (in the dehexanizer) into a C5/C6 streamthat is to be recycled to the MTP reactors andinto a heavier-hydrocarbons stream (gasoline).

The debutanizer overhead stream is mixedwith the vapor stream from the compressionstage and sent to the DME removal system. Thesystem overhead product, mainly propylene,is sent to the de-ethanizer, while the bottomis mostly recycled. The bottoms from the de-ethanizer is routed to the C3 splitter, to obtainpolymer-grade propylene, while the overheadis recycled to the MTP reactors. The bottomsmaterial from the C3 splitter is purged as liquid

petroleum gas (LPG).Economic performanceAn economic evaluation of the process wasconducted based on data from the fourth quar-ter of 2012. The following assumptions weretaken into consideration:• A 560,000 ton/yr unit erected on the U.S.

Gulf Coast (the process equipment is repre-sented in the simplified flowsheet)

• There is no storage for feedstock and product• Outside battery limits (OSBL) units consid-

ered: propylene refrigeration systemThe estimated capital investment (includingtotal fixed investment, working capital andother capital expenses) to build the MTP plant

is about $370 million. The capital investmentfor a similar plant built in China is about $295million. Figure 2 illustrates the operating costsfor both regions.

Currently, the existing MTP commercialunits were constructed in China, inside hugecomplexes, where coal is extracted andtransformed to synthesis gas, which is used tosynthesize methanol in high-capacity facilities(over 1 million ton/yr of methanol).

China possesses large reserves of coal,making it a favorable region for MTP plants.In the U.S., the growing availability of natural

gas extracted from shale guarantees theavailability of synthesis gas to produce themethanol required for the MTP process. There-fore, the country is also a promising region forfuture MTP units. ■



1. DME reactor

2. MTP reactors

3. Quench

4. Compression

5. Methanol recoverycolumn

6. DME removalsystem (col.1)

7. DME removalsystem (col.2)

8. Debutanizer column

9. Dehexanizer column

10. Deethanizer column

11. Propylene-propanesplitter

12. Refrigeration unit

CW Cooling water

RF Refrigerat ion fluid

ST SteamRF

MethanolCW

ST

LPG

ST

CW

ST

Water

CW

ST

RF

ST

CW

ST

Recycle 3

PG Propylene

CW

ST

Fuel gas

Recycle 1

Recycle 2

Recycles1, 2 and 3

CW

ST

CWST

CW

10

9

6

7

4

2

3

5

1

8

11

12

FIGURE 1. Propylene production from methanol, according to a process similar to the Lurgi MTP and JGC Mitsubishi DTP processes

0

200

400

600

800

1,000

1,200

1,400

1,600

1,800

U.S. GulfCoast

China

$ / m e t r i c t o n o f p r o d u c t

Raw materials costsMain utilities consumptionsFixed costs


FIGURE 2. Both China and the U.S. are favor-able regions for MTP plants because of coal in

China and shale gas in the U.S.



Polypropylene Production via Gas-

Phase Process: Stirred-bed reactor

Since 1956, polypropylene has beencommercially produced using catalystsand technology that were developed

from independent research by Karl Ziegler andGiulio Natta. Their namesake catalysts haveenabled the widespread use of this polymer inthe decades since. Polypropylene has becomea major part of the modern plastics-resinmarket because of its unique physical proper-ties, and is now employed in a wide range ofapplications from food packaging to automo-tive plastics.

Historically, polypropylene has beenproduced at industrial scale by three mainpolymerization approaches: hydrocarbonslurry or suspension processes, bulk-phaseprocesses and gas-phase processes. Twoleading technologies for gas-phase polypro-pylene production are Unipol, a fluidized-bedreactor process offered by the Dow Chemi-cal Co. (Midland, Mich.; www.dow.com),and Novolen, a stirred-bed reactor processdeveloped by Lummus Novolen, now a partof Chicago Bridge and Iron Co. N.V. (TheHague, The Netherlands; www.cbi.com). Aprocess similar to Unipol was discussed in thiscolumn last year (Chem. Eng., May 2013, p.33). Here, Figure 1 depicts a gas-phase pro-cess similar to Lummus Novolen’s stirred-bedreactor technology.

The processPropylene conversion to polypropylene is

achieved through the use of a Ziegler-Nattacatalyst. This process occurs in a continuousvertical stirred reactor at mild temperatures(about 80°C). The overall yields are typically>99 wt.%. The process can be divided in threemain areas: purification and reaction; resindegassing and pelletizing; and vent recovery.Purification and reaction. The polymerizationcatalysts are sensitive to several impurities,such as oxygen and water. To meet these

exposure restrictions, puri-fication equipment is usedbefore the reaction.

In the polymeriza-tion area, fresh purifiedpropylene is mixed with arecycled monomer streamand is fed to the reactoralong with the catalyst,co-catalyst and hydrogen.The heat of polymerizationis removed from the reactorby condensing propylenegas from the top of thereactor and recyclingthe liquid back into theagitated reactor.Resin degassing and pelletizing . The outlet reactorstream goes to a dischargevessel to be separated.The polymer powder is sent to a purge vessel,while the carrier gas goes through a cycloneand a filter and is finally sent for recovery.Nitrogen is used to remove residual monomerfrom the polymer powder and then the N2 gas is also sent for recovery.

The polymer powder then goes to a screwextruder, where the powder and additives aremelted, compounded and homogenized. Afterbeing extruded, the pellets are carried to acentrifugal dryer.Vent recovery. The carrier gas is compressed

and one portion of it is sent back to the reac-tor. The other portion is sent to a recoveryunit. The nitrogen and propylene streamsfrom the purge vessel are recovered andseparated in a membrane unit. The nitrogenis recovered to the process, while the mono-mer goes to a propylene-propane splitter(inside the purification area of the propylenesupplier, if the plant is part of a petrochemi-cal complex).

Economic performanceAn economic evaluation of the process wasconducted based on data from the first quarterof 2013 for a plant with a nominal capacityof 300,000 ton/yr, erected in the U.S. GulfCoast region (the required process equipmentis represented in the simplified flowsheet). Twoscenarios were analyzed:1) Integrated scenario. This case is based onthe construction of a plant linked to a propyl-ene supplier. In this scenario, the nearby unitcontinuously provides propylene, and receivesimpure propylene for purification. Thus, nostorage for propylene is required. However,storage of products is equal to 20 days ofoperation.2) Non-integrated scenario. This case cor-responds to a grassroots unit. Thus, 20 daysof operation was considered for both productsand raw materials. In addition, this scenarioincludes a propane-propylene splitter.

The level of integration is a key factor in thisprocess, since it can determinewhether the process is profitableor not. The level of integrationcauses the total fixed investment(Figure 2) and manufacturing

expenses to vary greatly. In termsof manufacturing expenses, innon-integrated plants, propyleneis purchased at higher prices,while in integrated complexes, itcan be obtained directly from thesupplier at a lower cost. n


Editor’s Note: The content for this col-umn is supplied by Intratec SolutionsLLC (Houston; www.intratec.us) andedited by Chemical Engineering. Theanalyses and models presented hereinare prepared on the basis of publiclyavailable and non-confidential infor-mation. The information and analysisare the opinions of Intratec and do notrepresent the point of view of any thirdparties. More information about themethodology for preparing this typeof analysis can be found, along withterms of use, at www.intratec.us/che.

Freshpropylene

(1) Reactor

(2) Fixed-bed dryers

(3) Discharge vessel

(4) Purge vessel

(5) Extruder

(6) Centrifugal dryer

(7) Compression

(8) Cooling tower

CW: Cooling water

Catalyst

Cocatalyst

H2

CW

N2

Additives

CWPropylenesupplier’s

purification

area (integratedpetrochemicalprocess)

Membrane

recovery

system

Offgasto flare

Polypropylene

pellets

1

2

3

4

6

5

7

CW 8

FIGURE 1. Polypropylene production via a gas-phase process similar to that of Lummus Novolen

0

50,000

100,000

150,000

200,000

250,000

300,000

Non-integratedscenario

Integratedscenario

U . S .

d o l l a r s ,

t h o u s a n d s Project

contingency

Indirect projectexpenses

Direct projectexpenses


FIGURE 2. Total fixed investment according to integration scenario



Polypropylene Production

via Gas-Phase Process

Polypropylene (PP) is one of the world’s most widely used poly-mers, second only to polyethylene in terms of global demand. Theglobal market for polypropylene is over 60 million metric tons per

year, and it is utilized in a broad and diverse range of end-uses — frominjection-molding applications to films and sheets, as well as syntheticraffia and other fibers, among others. Traditionally, the most representa-tive types of propylene polymerization are the following: hydrocarbonslurry or suspension, bulk (or bulk slurry), gas phase and hybrid (usesbulk- and gas-phase polymerization reactors).

The Unipol PP process, a leading gas-phase process technology, wasrecently offered for sale by Dow Chemical Co. (Midland, Mich.; www.dow.com). The company is looking to focus on high-margin areas, and isseeking buyers for its polypropylene licensing and catalyst business.

The processPP is a thermoplastic material formed by the polymerization of propyl-ene, resulting in a macromolecule that contains from 10,000 to 20,000monomer units. The production of a polypropylene homopolymer via a

gas-phase process similar to Dow Unipol is depicted in the diagram (Fig-ure 1). The process shown is capable of producting homopolymer andrandom copolymer PP. For impact copolymer production, a secondaryreaction loop is required. In this process, gaseous propylene contacts asolid catalyst in a fluidized-bed reactor. The process can be separatedinto three different areas: purification and reaction; resin degassing andpelletizing; and vent recovery.Purification and reaction. Fresh polymer grade (PG) propylene is sentto fixed-bed dryers to remove water and other polar impurities. Thepurified propylene, a recycle stream from the vent recovery system andcomonomers (in case of copolymer production) are then fed continu-ously to the reactor. A gas compressor circulates reaction gas upwardthrough the reactor, providing the agitation required for fluidization,backmixing and heat removal. No mechanical stirrers or agitators arerequired in the process reactor. The overhead gas from the reactor

passes through a cooler for reaction heat removal. Catalyst is continu-ously fed to the reactor.Resin degassing and pelletizing. Resulting granular polypropylene isremoved from the reactor by the discharge tanks and sent to a purgebin where residual hydrocarbons are stripped with nitrogen from theresin and are sent to the vent recovery system. The resin from purge binis combined with additives and then flows to the pelletizing unit. Thepellets are dried, cooled and sent to product blending and storage.Vent recovery. The vent gas is processed to separate hydrocarbons andnitrogen purge gas, which is returned to the process. The condensedcomponents are separated into a propylene stream, which is returned tothe reaction system, and a propane stream.

Economic performanceAn economic evaluation of the process was conducted for two distinct

integration scenarios:

• The integrated scenario is based on the construction of a plant linkedto a propylene supplier. In this case, storage for propylene is not re-quired. However, the estimated investment for the integrated scenarioincludes storage for PP equivalent to 20 days of plant operation

• The non-integrated scenario is based on the construction of a grass-roots PP processing unit. Thus, a time period of 20 days of operationwas considered for storage of both products and raw materials

The economic evaluation was based on data from the third quarter of2011 and a plant nominal capacity of 400,000 ton/yr erected on theU.S. Gulf Coast (only the process equipment is represented in the simpli-fied flowsheet).

The level of integration with nearby facilities significantly impactsthe total fixed investment required for the construction of a PP plant, asrepresented in the graph (Figure 2).

More than that, the elevated market prices for propylene in the U.S.make it unprofitable to operate a stand-alone PP unit. However, PP unitsthat are integrated upstream with a propylene production plant may pur-chase propylene at prices below the market average, reaching EBITDA(earnings before interest, taxes, depreciation and amortization) marginsof above 20%. ■


Editor’s Note: The content for this column is supplied by Intratec Solutions LLC(Houston; www.intratec.us) and edited by Chemical Engineering. The analysesand models presented herein are prepared on the basis of publicly available andnon-confidential information. The information and analysis are the opinions ofIntratec and do not represent the point of view of any third parties. More informa-tion about the methodology for preparing this type of analysis can be found, along

with terms of use, at www.intratec.us/che.

Integrated scenario

350,000

300,000

250,000

200,000

150,000

100,000

50,000

0

Indirectexpenses

Projectcontingency

Directexpenses

$ ,

t h o u s a n d s

Total fixed investment costs

Non-integrated scenario

1. Fixed-bed dryer

2. Polymerization reactor

3. Gas compression and cooling

4. Discharge and blowdown tanks

5. Purge bin

6. Horizontal ribbon blender

7. Extruder and pelletizer

8. Centrifugal dryer

9. P-P splitter

CW Cooling water

ST SteamPolypropylenepellets

Propanepurge

8

7

6

5

2

4

PG propylene

Catalyst andadditives

1

CW

3

Additives ST

9

CW


FIGURE 2. Elevated propylene prices make it unprofitable to operate astandalone PP unit

FIGURE 1. Homopolymer PP production process similar to Dow Unipol



Propylene Production

via Metathesis

Propylene is typically considered a co-product in steamcrackers and fluid-catalytic-cracking (FCC) processes,which are primarily driven by ethylene and motor gaso-

line production, respectively.Recently, most U.S. ethylene producers have shifted tolighter feedstocks due to the availability of low-cost naturalgas and natural-gas liquids (NGLs) from shale. This shiftdrastically decreases propylene byproduct output. Sincepropylene demand is growing faster than that of ethylene,and is expected to continue increasing, a gap between pro-pylene supply and demand is arising. This brings about thepossibility of establishing chemical processes for on-purposepropylene production, and one of the processes depends onmetathesis chemistry.

The metathesis processMetathesis, also known as disproportionation, is a reversiblereaction between ethylene and butenes in which carbon-

carbon double bonds are broken and then rearranged to formpropylene. Both ethylene and 2-butenes are mainly suppliedfrom steam-cracker units, but can also be obtained from FCCunits. The metathesis process depicted in Figure 1 is similar tothe OCT (olefins conversion technology) process developed byLummus Technology (part of Chicago Bridge & Iron Co. N.V.;the Hague, The Netherlands; www.cbi.com) and makes useof a tungsten oxide catalyst, along with a magnesium oxideco-catalyst. The OCT process typically achieves a propylene yield ofabout 90%.

The process is divided into two main areas: purification and reaction;and separation.Purification and reaction section. Fresh and recycled ethylene andbutene are mixed and fed to treatment equipment that removes potentialcatalyst poisons, such as oxygenates and sulfur. After treatment, the

stream is vaporized and heated to the reaction temperature, between280 and 320°C. The metathesis reaction occurs in a fixed-bed catalyticreactor, and butene is commonly employed in excess to minimize even-tual side reactions that produce mostly five- to eight-carbon olefins. Coke,another byproduct of the reaction, is deposited on the catalyst throughoutthe process. Each reactor can run for about 30 days before requiringregeneration, where coke is burned off in a controlled atmosphere.Separation section. Propylene purification is carried out in two col-umns. The stream leaving the reactor is cooled and sent to the deethylen-izer column, the overhead stream is recycled back to the reactor andthe bottom stream is fed to the depropylenizer column, which producespolymer-grade propylene in the overhead, as well as a heavies productstream (four-carbon compounds and greater) that is also recycled.

Economic performance

An economic evaluation of the process was conducted for three distinctlocations — the U.S. Gulf Coast region, Germany and Brazil — andis based on data from the third quarter of2011. The following assumptions were takeninto consideration:• A 350,000 ton/yr metathesis unit erected

inside a petrochemical complex (all equip-ment is represented in the simplifiedflowsheet in Figure 1)

• Storage cost of the main product is equalto 20 days of operation, and storage costfor feedstocks was not considered

• The excess fuel gas generated in the pro -cess is considered to be sold to a nearbychemical plant at natural gas prices

The estimated capital investment (includingtotal fixed investment, other capital expensesand working capital) and for such a planton the U.S. Gulf Coast is about $245 mil-

lion, the lowest among the regions compared. Germany presented ahigher capital investment — $295 million — but the lowest operatingcost, at about $1,320/ton, compared to $1,480/ton in the U.S. InBrazil, both the initial capital investment for a metathesis plant and theoperating costs were the highest.

Global perspective

Despite the higher operating costs, metathesis plants in the U.S. haveexhibited the highest EBITDA (earnings before interest, taxes, deprecia-tion and amortization) margins, since the propylene product prices inNorth America were higher than those in Europe.

The higher EBITDA, coupled with the lowest total capital investment,makes the U.S. the most promising region for new metathesis projects.However, the current number of metathesis units worldwide shows thatsuch units can also be profitable outside the U.S., as shown in the worldmap (Figure 2). ■


Editor’s Note: The content for this column is supplied by Intratec Solutions LLC(Houston; www.intratec.us) and edited by Chemical Engineering. The analyses andmodels presented herein are prepared on the basis of publicly available and non-confidential information. The information and analysis are the opinions of Intratecand do not represent the point of view of any third parties. More information aboutthe methodology for preparing this type of analysis can be found, along with termsof use, at www.intratec.us/che.

Fresh ethylene

Fresh butene

PGpropylene

Heavies to fuel

CW

ST

RF

ST

FU

FU

Lights to fuel

Fordisposal

1

5 6

23 3'

4

FU: fuel

CW: cooling water

RF: refrigeration fluid

ST: steam

(1) Reactor feed treater

(2) Reactor feed heater

(3) Metathesis reactor

(4) Regenerat ion gas heater

(5) Deethylenizer column

(6) Depropylenizer column

FIGURE 1.

Up to 149,000 ton/yr

From 150,000 to 199,000 ton/yr

From 200,000 to 249,000 ton/yr

At least 250,000 ton/yr

Each mark in themap correspondsto an existingmetathesis plant.The nominalcapacity of eachplant follows thelegend below:

FIGURE 2.




Propylene Production via

Propane Dehydrogenation

Because natural gas supplies are significantly increasing due tothe rising exploitation of shale gas, mainly in the U.S., propaneprices are decreasing. Coupled with low propane prices, ethyl-

ene producers are shifting to lighter feedstocks (more ethane, less naph-tha), which is decreasing yields of propylene in cracking operations.The increasing demand for propylene and the availability of low-costfeedstock make propane dehydrogenation an economically attractivechemical route.

Propane, the main feedstock for propane dehydrogenation (PDH)processes, can be obtained as a byproduct of petroleum refinery opera-tions and can be recovered from propane-rich liquefied petroleum gas(LPG) streams from natural-gas processing plants.

The PDH processPDH is an endothermic equilibrium reaction. The PDH process depictedbelow is similar to the Oleflex process developed by UOP LLC (DesPlaines, Ill.; www.uop.com), and is suited to produce polymer-grade(PG) propylene from propane. The maximum unit capacity is around

650,000 ton/yr. This process is carried out in the presence of a plati-num catalyst and achieves overall propylene yields of about 90 wt.%.The industrial plant can be divided into two main sections: reaction

and product recovery.Reaction. In the reaction section, after heavy impurities removal inthe de-oiler column, propane is sent to the dehydrogenation reactors.The propylene yield in such reactors is favored by higher temperaturesand lower pressures. However, temperatures that are too elevated willpromote thermal cracking reactions that generate undesirable byprod-ucts. Therefore, the PDH reaction usually occurs at temperatures of about650°C and near atmospheric pressures.

In order to purge the coke accumulated on the catalyst surface duringthe reaction, a continuous catalyst regenerator (CCR) unit is required. Thecatalyst circulates in moving beds through the reactors, before being fedto the CCR unit, which operates independently of the reaction, burning off

the coke and returning the catalyst to its reduced state.Product recovery. The reactor effluent is compressed, dried and sent tothe product recovery section. In this section, a hydrogen-rich stream isrecovered and light hydrocarbons and hydrogen traces are removedin a de-ethanizer. The PG propylene product is further purified in apropane-propylene (P–P) splitter and leaves as the top product.

Economic performanceAn economic evaluation of the process was conducted based on datafrom the second quarter of 2012. The following assumptions are as-sumed for the analysis:• A 550,000 ton/yr PDH unit erected inside a petrochemical complex

(all equipment represented in the simplified flowsheet below)• No storage of feedstock and product is considered• Net raw materials cost is the difference between propane and catalyst

make-up costs and credits from fuel and electricity generated in theprocess

Global perspectiveRecently, conditions in the U.S. have led to the lowest production costsand the most attractive EBITDA (earnings before interest, taxes, depre-ciation and amortization) margins (about 30%), due to the availability

of low-cost propane derived from shale gas. Low-cost propane importedfrom the Middle East allows China to present favorable EBITDA marginsof about 20%. The optimism about this process is demonstrated by therecently announced plans for at least six PDH units in China.

Considering the capital-cost requirements presented and an operatingrate of 91%, the internal rate of return is above 30% in the U.S. andabout 20% in China.

On the other hand, South America and Europe do not offer favorableconditions for PDH units. ■


Editor’s Note:The content for this column is supplied by Intratec Solutions LLC. (Houston; www.intratec.us) and edited by Chemical Engineering. The analyses and models pre-sented herein are prepared on the basis of publicly available and non-confidentialinformation. The information and analysis are the opinions of Intratec and do notrepresent the point of view of any third parties. More information about the meth-

odology for preparing this type of analysis can be found, along with terms of use,at www.intratec.us/che.

0

200

400

600

800

1,000

1,2001,400

US Gulf Brazil China

Regional cost comparison

Net rawmaterials costs

Main utilitiesconsumptions

Fixedcosts

200

300

400

500

600

700

US Gulf Brazil China

$ , m i l l i o n s

$ / m i l l i o n t o n p r o d u c t

Total capital investment

Total fixedinvestment (TFI)

Workingcapital

Other

Catalyst flow

(1) Reactors

(2) Reactor heaters

(3) CCR

(4) Compress ion & drying

(5) Cold box

(6) PSA

(7) Selective hydrogenation

(8) Dee thanizer

(9) P-P Splitter

(10) De oiler

(11) C3= Refrigeration unitFresh propane

PG propylene

Heaviesto fuel

Lights to fuel

1

3

4

5

7

8

9

6

ST

RF

RF: Refrigeration fluid

112

10

FU

CW

CCR: continuous catalyst regeneration PSA: pressure-swing absorption

: Cooling water ST: SteamFU: Fuel

CW

RF




Sodium Hypochlorite

Chemical Production

Sodium hypochlorite (NaClO) is the active constitu-ent in chlorine bleach, a strong oxidizer andbleaching agent. Increases in household bleach

demand are driven mostly by population growth. Inturn, population growth and its corresponding increasesin water consumption — coupled with limited freshwater resources — makes water treatment the largestapplication for bleach, as well as the fastest-growingsegment of bleach use.

Sodium hypochlorite chemical production is a well-established process in the industry, and the principlebehind its operation is also employed for preventingchlorine emissions in chlor-alkali plants. The chemicalprocess relies on the acquisition of chlorine and causticsoda (sodium hydroxide; NaOH) feedstock from exter-nal sources, in contrast with the electrochemical processfor bleach, which also involves brine electrolysis.

The process

The chlorination of caustic soda to sodium hypochlorite is an exother-mic reaction. The chemical production process depicted below is awidely used process, similar to the one employed by Solvay Chemicals(Brussels, Belgium; www.solvay.com), for example. The process is suitedto producing both household bleach (5–6 wt.%) and industrial bleach(10–15 wt.%), and relies on the chlorination of caustic soda by chlorinegas within packed columns.

The industrial sodium-hypochlorite plant can be divided into two mainsections: reaction and product discharge; and bleach filtration.Reaction and product discharge. The chlorine absorption system canbe divided into two parts: in the first, a packed column is operatedwith a safe excess of caustic to prevent reduction in pH; the second

part, in turn, receives the liquor from the first column to be post-chlorinated, until the desired bleach concentration is reached. Priorto reaction, caustic soda is diluted with water in the first buffer tank,along with the first-column bottom stream. Chlorine gas is dilutedwith air and fed into the bottom of both columns. Flow to the secondpart of the system is established when the finished product is sent tothe filtration steps.Bleach filtration. Bleach filtration is necessary to meet product quality re-quirements and is often the last step before storage. Usually, the bleachfilter system consists of pressure-leaf filters. After filtration, the product issent to storage. Backwash water containing spent filter aid can be takento a dry-cake pressure filter for further processing.

Sodium hypochlorite solutions are very sensitive and special opera-tions must be carried out to prevent its decomposition. Besides pH,other factors affecting degradation are the initial bleach concentration,exposure to ultraviolet light, presence of certain metals and elevated

temperatures. The graph above shows temperature dependency forbleach degradation, starting from a 12.5 trade-percent solution. Tradepercent is an expression of available chlorine, in units of grams per literof available chlorine.


An economic evaluation of the process was conducted for three distinctlocations — the U.S. Gulf Coast region, Germany and Brazil — and isbased on data from the first quarter of 2012. The following assumptionsare made for the analysis:• A 250,000 ton/yr chemical production unit (NaClO solution, 12.5

wt. %) erected inside a chlor-alkali facility (all equipment is repre-

sented in the simplified flowsheet)• Storage of products is equal to 20 days of operation, and there is nostorage for feedstock

The estimated capital investment for such a plant on the U.S. Gulf Coastis about $35 million, the lowest among the regions compared. Germanypresented a higher capital investment, at $40 million, and the highestoperating costs — about $170/ton (compared to $140/ton in the U.S.)

Global perspective

Selling prices for bleach vary substantially, depending on supply anddemand uctuations of the chlorine/caustic market. Based on theestimated capital and operating costs, a U.S. Gulf Coast-based bleachmanufacturing venture can reach an internal rate of return (IRR) above15%, by selling bleach at an average price of $175/ton, a reasonablevalue. However, a venture in Germany would not be economically at-tractive at the same selling price. ■


Editor’s Note: The contentfor this column is suppliedby Intratec Solutions LLC(Houston; www.intratec.us) and edited by ChemicalEngineering. The analysesand models presented hereinare prepared on the basis ofpublicly available and non-confidential information. Theinformation and analysis arethe opinions of Intratec anddo not represent the pointof view of any third parties.

More information about themethodology for preparingthis type of analysis can befound, along with terms ofuse, at www.intratec.us/che.

5

6

7

8

9

10

11

12

13

0 20 40 60 80 100 120

B l e a c h c o n c e n t r a t i o n ,

t r a

d e %

Time, days

15 °C 20 °C 25 °C 30 °C 40 °C

Caustic soda

Offgas toscrubber

Water

Chlorine

Bleachto storage

Backwash

water todisposal

Air

2 4

CWCW

1 3

5

6

(1) 1st Absorptioncolumn

(2) 1st Buffertank

(3) 2nd Absorptioncolumn

(4) 2nd Buffer

tank

(5) Bleach filtersystem

(6) Backwash

water filter

CW = Coolingwater




n -Butanol from

Lignocellulosic Feedstock

The alcohol n-butanol is applied in the production of butyl acrylatesto supply the coatings and adhesives industries. It is also used in

plasticizers or directly as a solvent. It is commercially producedfrom propylene and synthesis gas (syngas) by chemical synthesis.Due to the increasing interest in biofuels development, the produc-

tion of n-butanol through a fermentation process is being revisitedand improved. Biobutanol is an attractive renewable fuel that can beproduced from the fermentation of sugars derived from agriculturalfeedstock — for example, corn and sugarcane — or from waste ligno-cellulosic materials, such as corn stover and sugarcane bagasse.

In this field, bio-butanol is viewed as an alternative to ethanol, sinceit fits the existing fuel infrastructure better and exhibits higher energycontent, which makes it a more suitable gasoline blending fuel.

The processThe production process for n-butanol from corn stover via acetone-butanol-ethanol (ABE) fermentation is depicted in Figure 1. The feed

handling, pretreatment and fermentation steps were compiled from atechnical report published by the National Renewable Energy Labora-tory (Golden, Colo.; www.nrel.gov; NREL/TP-5100-47764, May 2011)and adapted for n-butanol production. The purification step was basedon typical ABE fermentation processes.Feed handling. Corn-stover supply trucks send the biomass to weightingand unloading stations, followed by a short-term queuing storage andconveyors for feeding the feedstock to the pretreatment area.Pretreatment. The biomass, composed of cellulose, hemicellulose andlignin, is treated with dilute sulfuric acid and heated in a screw-feedreactor to convert most of the hemicellulose into fermentable sugar, suchas xylose. After pretreatment, the solution is flash-cooled, and water isremoved, condensed and sent to the wastewater-treatment area. The mix-ture is then diluted with water and its pH adjusted by adding ammonia.Hydrolysis and fermentation. In this step, cellulose is converted to

fermentable glucose using enzymes. The hydrolysis is initiated in acontinuous reactor and completed in several parallel batch fermenters.The slurry is then cooled and inoculated with microorganisms, whichconvert xylose and glucose into acetone, ethanol and n-butanol. Thefermentation beer is then sent to the beer well before being directed tothe purification section.Purification. The fermentation beer is separated into acetone, ethanoland n-butanol products by a series of distillation columns, a decanterand a molecular sieve unit (to purify ethanol to commercial grade). Thebottoms of the beer column, containing water and lignin, is sent to afilter. The filtrate is directed to the wastewater-treatment apparatus. Theresidues from wastewater treatment and lignin from filtration are burnedfor steam and electricity generation. The distillate from the n-butanol col-umn is recycled to the acetone column to improve n-butanol recovery.


An economic evaluation of the process was conducted taking intoconsideration a facility located on the U.S. Gulf Coast that is capableof producing 100,000 ton/yr of bio-butanol, 10,000 ton/yr of acetoneand 8,000 ton/yr of ethanol.

The total estimated investment (including total fixed investment, work-ing capital and initial expenses) for this plant is about $450 million,while the operating expenses are about $1,000/ton of n-butanol.

Global perspectiveIn the past, the ABE fermentation was used to commercially produceacetone and n-butanol, but this process fell out of favor because petro-chemical processes became more cost-effective.

Still, the production of bio-butanol from ABE fermentation presentsseveral challenges, such as low n-butanol yield, concentration and pro-ductivity, when compared to ethanol biofuel. However, recent research

has been focused on enhancing such issues by developing improvedmicroorganisms, fermentation techniques and low-energy recoverymethods, in order to reduce production costs.

A good strategy to prove the bio-butanol production technology atcommercial scale is to target the chemicals market, which offers highermargins and a broad range of applications (Figure 2). Later, consid-eration could be given to entering the fuels market, which offers farhigher volumes, but lower margins. n

Editor’s Note: The content for this column is supplied by Intratec Solutions LLC(Houston; www.intratec.us) and edited by Chemical Engineering. The analysesand models presented are prepared on the basis of publicly available andnon-confidential information. The content represents the opinions of Intratec only.More information about the methodology for preparing analysis can be found,along with terms of use, at www.intratec.us/che.

1) Feed handling2) Screw-feed reactor3) Flash vessel

4) Ammonia conditioning5) Continuous hydrolysis6) Batch hydrolysis/fermentation

7) Beer well8) Beer column9) Acetone column

10) Ethanol column11) Molecular sieve unit12) Decanter

13) n-butanol column14) Pressure filter15) Wastewater treatment unit

16) Energy cogeneration unit17) Cooling tower

BFW boiler feed waterCW cooling waterEE electricityST steam

2

Corn Stover 1

34

5

6 7

15

14 16

17

8 9 10 13

12

Sulfuric acid

ST

CW

ST

BFW

Ammonia

ST ST ST

Enzyme CW

CW CW CW

Acetone

Water

Ethanol

Lignin ST

EE

CW

n-Butanol

11

FiltrateResidues

CW

FIGURE 1. The ABE fermentation production process produces n-butanol

n Acrylate 32%n Exports 26%

n Glycol ethers 20%

n Acetate 12%

n Solvent 5%

n Plasticizer 2%

n Other 2%

n Amino resins 1%

FIGURE 2. Recent U.S. n-butanol demand



Sugar Inversion

The substance known colloquially as table sugar is actually su-crose, a disaccharide composed of fructose and glucose. Sucrose

is a major agriculturally derived product with an established in-dustry for its extraction, processing and supply. It is mainly employed asa sweetener in industrial and domestic food applications, but also findsadditional uses in the pharmaceutical industry, in chemical manufactur-ing, and as a feedstock for fermentation processes. For example, sugaris used as a feedstock for the manufacture of bio-based chemicals, suchas bio-succinic acid.

Sugar generally exists commercially as a solid granular product.However, some food manufacturers prefer to use sugar in a liquid form,due to the ease of handling a liquid product. One commercially avail-able liquid-sugar product is l iquid invert sugar, also known as invertsyrup. It is produced from the inversion of sucrose, which refers to thehydrolysis of the disaccharide molecule into its constituent parts, themonosaccharides glucose and fructose. The result of this reaction is aproduct with greater sweetening power and improved microbiological

stability, when compared to sucrose.Invert syrups are commercialized with different combinations of invertsugar and sucrose contents, depending on the degree of inversion per-formed. Also, invert syrups can be produced by three different inversionprocesses: acid hydrolysis with mineral acids; enzymatic hydrolysis; orhydrolysis by cation ion-exchange resin. The latter process is describedin this column.

The processThe process for sucrose inversion by ion-exchange resin shown inFigure 1 is similar to the one presented in U.S. Patent 8,404,109, pub-lished by European Sugar Holdings S.a.r.l. (Capellen, Luxembourg). Animportant feature of this process is the removal of ash from the solutionin the ion-exchange columns.Sucrose dissolution. Water and steam are mixed to form a hot water

stream. Part of this stream is added to raw sugar (sucrose) before it isfed to a screw conveyor, which directs wet sucrose to an agitated ves-sel. The remaining hot water is fed into the vessel, forming a 60 wt.%sucrose solution.Sucrose inversion. The sucrose solution is fed to a cation-exchangecolumn, where some ash is retained and the pH is lowered, allowingthe inversion reaction to occur. Then, part of the inverted solution is fedto an anion-exchange column, where some ash is removed and the pHof the solution is increased. Finally, the fully inverted solution is concen-trated to 70 wt.% dry matter, forming the final product.

Economic evaluationAn economic evaluation of the sugar-inversion process was conducted,taking the following assumptions into consideration:

A facility producing 585,000 ton/yr of invert syrup from raw sugar.•

The facility was assumed to be located on the U.S. Gulf Coast in thesame location as a plant operating a fermentative process that usesinvert syrup as feedstock

A raw-sugar warehouse with storage capacity equal to 20 days of•

operation Utilities facilities and product storage were not considered in the•

analysis

The estimated total fixed investment for building such a plant is about$40 million.

Global perspective

Sucrose can be produced either from sugarcane or sugarbeet plants,depending on climate conditions. Sugarcane is cultivated in tropicaland subtropical regions, while sugarbeets are more suitable in temper-ate zones.

In 2013, world sugar production was about 177 million tons (in rawsugar equivalent). About 80% of global sugar production is derivedfrom sugarcane, and about 20% is from sugarbeets.

Brazil, the world’s largest sugar producer, accounts for about 22%of global sugar production, with all of its sugar production originatingfrom sugarcane.

The world’s main sugar producers are shown in Figure 2. The U.S. isthe only country among them that produces sugar from both sugarcaneand sugarbeet in significant amounts. As the world’s sixth-largest sugarproducer, the U.S. produces 56% of its sugar from sugarbeets, whilethe remaining 44% comes from sugarcane. n

Editor’s Note: The content for this column is supplied by Intratec Solutions LLC(Houston; www.intratec.us) and edited by Chemical Engineering. The analysesand models presented are prepared on the basis of publicly available andnon-confidential information. The content represents the opinions of Intratec only.More information about the methodology for preparing analysis can be found,along with terms of use, at www.intratec.us/che.

1) Screw conveyor2) Dissolution vessel3) Cation exchange resin bed

4) Anion exchange resin bed5) Evaporator

CW Cooling waterST Steam

1

2

3 4

5

H2O

ST

Raw Sugar

CW

CW

ST

To wastetreatment

Liquidinvertsugar

FIGURE 1. The sugar-inversion process shown here uses ion-exchange resins

n United States 5%

n Brazil 22%

n European Union 9%n India 15%

n China 8%

n Thailand 6%

n Mexico 4%

n Others 31%

FIGURE 2. World sugar production by country or region in 2013


http://slidepdf.com/reader/full/technology-profiles-chemical-engineering 22/31CHEMICAL ENGINEERING WWW.CHEMENGONLINE.COM FEBRUARY 2015

Technology Profile

39

‘Green’ Ethylene Production


W

ith a global nominal ca-pacity of about 140 millionton/yr, ethylene is amongthe main petrochemicals

produced worldwide, and is a keybuilding block for the industry. Eth-ylene is produced mostly via steam-cracking of petroleum-based feed-stocks, such as naphtha. Ethylene isa raw material for the manufacture ofpolyethylene, polyvinyl chloride (PVC),ethylene oxide and other products.

Global concerns about sustainabilityand global warming have promptedthe chemical industry to develop pro-duction routes for ethylene that utilize

non-petroleum resources. Renewableethylene-production alternatives havebegun to emerge in this context.

“Green” ethylene can be producedby the dehydration of ethanol, whichcan be produced from renewable feed-stocks such as sugarcane and corn.Ethanol-derived ethylene is chemi-cally identical to traditional ethylene, sodownstream processing is equivalent.

The process

A process for ethylene productionvia ethanol dehydration similar to theprocesses developed by Chematur Technologies AB (Karlskoga, Swe-den; www.chematur.se) and PetronScientech Inc. (Princeton, N.J.; www.petronscientech.com) is depicted inFigure 1. The reaction occurs in fourfixed-catalyst-bed adiabatic reactors. Reaction. Ethanol is vaporized,heated in a furnace and fed to thefirst reactor. During reaction, the tem-perature drops and the output streammust be re-heated before entering the

next reactor. This is repeated until thefourth reactor, where ethanol reaches99% conversion. Next, the reactor-product stream is cooled in a waste-heat boiler, generating steam.Quenching, compression, causticwashing and drying. After reaction,the product stream’s water contentis reduced in a quench column. Theethylene stream leaves by the over-heads and is compressed beforeentering the washing column, whereimpurities are removed with a causticsolution. Process water is also sup-plied to this column to avoid causticentrainment along with the ethylene

overhead stream, which is cooled byinterchange with the ethylene prod-uct stream. Water separated in thecompression and cooling stages isrecycled to the quench column. Next,the gas stream is dehydrated in a mo-lecular sieve unit. Purification. After cooling by inter-change with the ethylene product, thedehydrated stream is fed into the eth-ylene column, where heavy residuesare removed. Next, light impurities are

removed in the stripper column. Thetwo columns share a single condenser,which uses propylene refrigerant. Thepolymer-grade ethylene (99.9 wt.%)product is obtained in the bottoms ofthe stripper column and used to coolprocess streams in heat integrations.

Economic performance An economic evaluation of the pro-cess assumes the following:

A facility capable of producing•

300,000 ton/yr of green ethyl-ene constructed on the U.S. Gulf

Coast, partially integrated with apolyethylene complexStorage capacity equal to 20 days•

of operation for ethanol and nostorage for ethylene

The estimated capital expenses (to-tal fixed investment, working capitaland initial expenses) to construct the

plant are about $260 million, while theoperating expenses are estimated atabout $1,400/ton of green ethylene.

Global perspective The cost of manufacturing ethylenefrom ethanol dehydration is highlydependent on raw material cost (Fig-ure 2). Because ethanol is used as abiofuel and the crop feedstocks usedin its production are also part of thefood market, ethanol prices may vary

according to the market dynamics ofboth food and biofuels.However, a niche consumer market

may pay higher prices to obtain envi-ronmentally friendly products. There-fore, it is possible to commercializegreen ethylene at a premium price,sufficient to offset the higher produc-tion costs of green ethylene. n

Editor’s Note: The content for this column is supplied by IntratecSolutions LLC (Houston; www.intratec.us) and edited by ChemicalEngineering . The analyses and models presented are prepared on thebasis of publicly available and non-confidential information. The contentrepresents the opinions of Intratec only. More information about the

methodology for preparing analysis can be found, along with terms ofuse, at www.intratec.us/che.

1) Ethanol vaporizer2) Furnace3) Reactors4) Quench column5) Compression system6) Caustic washing column7) Molecular sieve unit8) Ethylene column9) Stripper column10) Cooling tower11) Refrigerant system12) Steam boiler

BFW Boiler feed waterC2= Ethylene productCW Cooling water

FU FuelPW Process waterRF RefrigerantST Steam

1

2

4

5

EthanolST

FU

BFW

CW

To wastetreatment

CW

PW

C2=

Causticsoda

To wastetreatment

C2=

ST

RF

Lightsto fuel

ST

Heaviesto fuel

CW

RF

ST

6

7

8 9

10

12

11

3

Ethylene product(C2=) to heatintegrations

FIGURE 1. Green ethylene production via an ethanol dehydration process that is similar to Chematur and Petron processes

n Net raw materialscost 89%

n Main utilitiesconsumptions 6%

n Fixed costs 5%

FIGURE 2. Operating expenses for green ethylene


http://slidepdf.com/reader/full/technology-profiles-chemical-engineering 23/31CHEMICAL ENGINEERING WWW.CHEMENGONLINE.COM MARCH 2015

Technology Profile

48

Propane Dehydrogenation: Oxydehydrogenation


P

ropylene is a major compo-nent of the global olefins mar-ket and is widely used as anintermediate for an array of

chemical and plastic products.Propylene is produced as a byprod-

uct in steam crackers and throughfluid catalytic cracking (FCC) process-es. However, due to the increasedavailability of inexpensive ethane fromshale gas, ethane has been prefer-entially used as a feedstock in steamcrackers over naphtha, a practice thatresults in minor propylene production.

In this context, on-purpose propyl-ene production routes are of great

interest to the petrochemical market-place. Among them, propane dehy-drogenation (PDH) technology is an at-tractive option, and has been appliedin several plants around the world.

The process The PDH process depicted in Figure1 is similar to the STAR (steam activereforming) process (Uhde GmbH; Dort-mand, Germany; now ThyssenKruppIndustrial Solutions; www.thyssenk-

rupp.com), which applies the oxydehy-drogenation principle. Two other PDHprocesses were previously discussedin this column (Chem. Eng., Feb. 2013,p. 33 and Jan. 2014, p. 27).

The STAR process was the first toapply the oxydehydrogenation principlefor propylene production, significantlyenhancing performance. In this tech-nology, O2 is added to the system toreact with H2, forming water. This reac-tion lowers H2 partial pressure, shiftingthe equilibrium to higher conversion ofpropane into propylene while providing

the required heat of reaction. Figure 2shows the effect of O2 addition on pro-pane conversion, according to differentO

2-to-propane molar ratios.

Reaction. Propane is fed into a depro-panizer column for heavy impuritiesremoval. The depropanizer distillate isvaporized, mixed with steam, and pre-heated before entering the reformer re-actor, which is an externally fired tubu-lar reactor where most of the propaneconversion occurs. Next, the reformereffluent is fed to the oxyreactor, whereH2 is selectively combusted by O2. The process gas is cooled in a seriesof heat exchangers to recover process

heat. Then, it is compressed and sentto the next area.CO 2 separation. The compressedgas is fed to an absorption column,where CO2 is removed in the bottomsby washing with an aqueous pipera-zine-activated methyldiethanolamine(MDEA) solution. The bottoms streamis depressurized, liberating gases thatare recycled to the absorption column. The CO2-rich liquid solution from theflash vessel is fed to the solvent re-

generation column, which separatesCO2 in the overheads from the MDEAsolution in the bottoms, which is re-cycled to the absorption column.Gas separation. The overheadstream of the absorption column isdried in a molecular sieve and fed toa low-temperature separation system(cold box) for the removal of lights. Inthis system, a H2-rich stream is alsoobtained, and is purified in a pressureswing adsorber (PSA) unit to obtainhigh-purity H

2 byproduct.

Fractionation. The liquid fraction

from the cold box is sent to the frac-tionation area, which consists of a de-

ethanizer column for removal of lightsand a propane-propylene (P-P) splitter,where polymer-grade (PG) propylene isobtained. Unreacted propane from thesplitter bottoms is recycled.

Economic performance An economic evaluation of a PDH plantwas conducted, assuming a facilitywith a nominal capacity of 450,000ton/yr of PG propylene erected at apetrochemical complex on the U.S.

Gulf Coast (no storage for feedstockand product was considered).Estimated capital expenses (total

fixed investment, working capital andinitial expenses) for such a plant are$400 million and estimated operatingexpenses are $770/ton of product.n

Editor’s Note: The content for this column is supplied byIntratec Solutions LLC (Houston; www.intratec.us) and editedby Chemical Engineering . The analyses and models pre-sented are prepared on the basis of publicly available andnon-confidential information. The content represents the opin-ions of Intratec only. More information about the methodologyfor preparing analysis can be found, along with terms of use,at www.intratec.us/che.

1) Depropanizer column2) Reformer reactor3) Feed pre-heater4) Oxyreactor5) Compression system6) Absorption column7) Flash vessel8) Regeneration column9) Molecular sieve unit10) Cold box11) PSA unit12) Deethanizer column13) P-P splitter14) Refrigeration unitMDEA Methyl

diethanolamine

BFW Boiler feed waterCW Cooling waterFU FuelRF Refrigeration fluidST Steam

Propane

ST

FU

O2 /air

CW

CW

BFW

H2 byproduct RF

RF

ST

Air

ST

ST

BFW

ST

ST

ST

Watermake-up

MDEAmake-up

Piperazinemake-up

CO2

RF

Lightsto fuel

Lightsto fuel

CW

PGpropylene

1

2

45

6

7

8

9

10

12

11

3

13

14

Heaviesto fuel

FIGURE 1. Propane dehydrogenation process similar to the Uhde (now ThyssenKrupp) STAR process

FIGURE 2. O2 impact on propane conversion

60

55

50

45

40

35

30

25

20

Without oxygen

02 /HC ratio = 0.05

02 /HC ratio = 0.1

550 560 570 580 590 600

Temperature (oC)

C o n v e r s i o n p r o p a n

e ( % )



Technology Profile

CHEMICAL ENGINEERING WWW.CHEMENGONLINE.COM APRIL 201536

Onsite Enzyme Production


E

nzymes are proteins synthe-sized in living cells that actas biocatalysts for specificreactions. However, enzymes

must be applied under mild condi-tions, since they lose their catalytic ac-tivity when subjected to heat, organicsolvents or strong acids and bases.

Enzymes are industrially producedfrom the fermentation of microorgan-isms, but they can also be obtainedfrom plant and animal sources. Theyoffer high specificity and selectivityfor reactions and substrates, and canproduce enantiopure compounds (incontrast to chemical syntheses, which

usually produce racemic mixtures).Enzymes are applied in the produc-

tion of food, fine chemicals and phar-maceuticals. In addition, they are usedin washing processes, as well as fora wide range of analytical purposes. Alternatively, enzyme technology hasbeen developed to improve sustain-able processes that produce existingproducts using novel raw materials,such as biomass.

The process A typical process for onsite enzymeproduction through fermentation is de-scribed below, using the enzyme cel-lulase as an example. Cellulase is usedin the hydrolysis of lignocellulosic mate-rial in biomass-based processes, suchas the production of cellulosic ethanol.Cellulase is industrially produced by afungus in an aerobic fermentation inthe presence of a cellulase-productioninducer (Figure 1). The process shownwas compiled from a technical reportpublished by the National Renewable

Energy Laboratory (NREL; Golden,Colo.; www.nrel.gov). Medium preparation. Glucose feed-stock is mixed with water and a smallquantity of cellulase, which convertspart of the glucose to sophorose, a

cellulase-production inducer. Pre-fermentation. The fermentationinoculum is prepared in a batch pre-fermentation step, which occurs infour trains with three pre-fermenterseach, into which part of the glucosemedium, nutrients and air are fed. Fermentation. The fermentation stepis conducted in nine fermenters, whichare fed with the fermentation inoculum,culture medium, nutrients and air. Thefermentation is a fed-batch process, in

which, after the depletion of nutrientscaused by cell growth, glucose me-dium is fed to the reactors, allowingcontinuous enzyme production. Afterfermentation, the cellulase-containingbroth can be directly fed to the hydroly-sis section of a cellulosic ethanol facil-ity, since the high temperature appliedin this operation is sufficient to inacti-vate the microorganism.

Economic evaluation An economic evaluation of the pro-cess was conducted based on data

from Q2 2014. The following assump-tions were considered:

A cellulase production unit con-•

structed in the same complex as acellulosic ethanol facility located onthe U.S. Gulf Coast

Onsite enzyme production unit man-•

ufactures 5,200 ton/yr of cellulaseStorage and utility facilities were not•

considered Total fixed investment for the con-

struction of this unit was estimated atabout $40 million.

Global perspectiveCellulase enzyme is applied to hydro-lyze lignocellulosic biomass, which isused as feedstock in the production

of cellulosic ethanol (Figure 2). In thisway, cellulosic ethanol costs are highlyrelated to enzyme costs.

Initially, enzyme costs were too high,so cellulosic ethanol was not cost-competitive with corn-based ethanol.Currently, many advances in enzymetechnology have been achieved, low-ering enzyme costs in cellulosic etha-nol by both enzyme production-costreduction and enzyme efficiency en-hancement. In addition, onsite pro-duction allows the manufacture of thedesired enzyme at lower costs, sincesome downstream operations presentin a standalone enzyme unit (such aspurification, formulation and storage),as well as transportation of the en-zyme product, are not necessary. n

For more on industrial enzymes, see “Tunable En-zymes and the Leaner, Greener CPI” in the January

2015 issue of Chem. Eng . at www.chemengonline.com

Editor’s Note: The content for this column is supplied by In-tratec Solutions LLC (Houston; www.intratec.us) and edited byChemical Engineering. The analyses and models presented are

prepared on the basis of publicly available and non-confidentialinformation. The content represents the opinions of Intratec only.More information about the methodology for preparing analysiscan be found, along with terms of use, at www.intratec.us/che.

1) Medium tank 2) Nutrients tank

3) Pre-fermentationcompressor

4) Pre-fermenters5) Fermentation compressor

6) Fermenters

CW Cooling water

Glucose

water

AirCW

1

2

4

5

6

3 CW

Nutrients

Cellulase

Cellulasebroth

Air

FIGURE 1. The process diagram shows onsite cellulase enzyme production

FIGURE 2. A typical cellulosic ethanol production process is shown here

Pretreat-ment

Enzymatichydrolysis

Fermenta-tion Distillation

Cellulosicethanol

Lignocellulosic

biomass

Cellulase

Solid residues



Technology Profile

CHEMICAL ENGINEERING WWW.CHEMENGONLINE.COM MAY 201548

Hydrogen Production from Natural Gas


H

ydrogen (H2 ) is an importantchemical feedstock, mainlyapplied in the manufactureof ammonia and methanol,

and for hydroprocessing operations inpetroleum refineries. Also, since H2 isan energy carrier, it has been consid-ered for stationary power and trans-portation applications.

Hydrogen production technolo-gies are separated into three maincategories: thermal, electrolytic andphotolytic. In thermal processes, suchas reforming and gasification, H2 isproduced from biomass and fossil fu-els, such as coal and natural gas. In

electrolytic processes, H2 is obtainedfrom water-splitting, using electricitythat can be generated from a varietyof sources, such as wind. In photolyticprocesses, light energy allows hydro-gen production using novel photo-electrochemical and photobiologicalwater-splitting processes.

In the U.S., H2 is mostly producedfrom natural gas using the thermalsteam methane reforming (SMR)process. Natural gas is an important

feedstock for H2 production since it iswidely available and presents a highhydrogen-to-carbon ratio, reduc-ing the generation of carbon dioxide(CO2 ) byproduct.

The processIn the process described below anddepicted in Figure 1, H2 is producedfrom natural gas using an SMR pro-cess. The process was compiledbased on information available in thechemical literature.Sulfur removal. Natural gas feed-

stock is purified by catalytic treatmentwith H2 for removal of sulfur impurities.In the hydrotreater, H2 reacts, over acatalyst, with sulfur compounds pres-

ent in the feed stream to form hydro-gen sulfide (H2S), which is then ad-sorbed in the desulfurizer.Steam reforming. Purified naturalgas is mixed with high-temperaturesteam and reformed into CO and H2. The reforming reaction requires a largeamount of heat and takes place in anexternally fired tubular reactor filledwith catalyst.Water-gas shift. CO and steam reactin a catalytic water-gas shift reaction,

forming additional H2 and CO2. Purification. CO2 and other impuritiesare removed from the H2 stream in apressure-swing adsorption (PSA) sys-tem. The purge stream from the PSAsystem is recycled to the reformer,where it is burned with fuel to provideheat to the reaction. The H2 productobtained has purities of 99.99 wt.%.

Economic evaluation An economic evaluation of the pro-cess was conducted based on thefollowing assumptions:

A central facility with a nominal•

capacity of 450,000 ton/yr of H2 erected on the U.S. Gulf CoastDistribution costs and storage for•

feedstock and product were notconsidered The estimated total fixed invest-

ment for the construction of this plantis about $460 million.

Global perspective There are three kinds of facilities for H2 production: central, semi-central anddistributed facilities (Figure 2). Theydiffer in their location and scale of pro-duction, characteristics that directly

affect H2 cost, competitiveness andtimeframe to market.Central facilities are located far from

the H2 point of use and are able to pro-duce large amounts of H2, benefitingfrom economies of scale. This type offacility requires high capital investment,as well as a distribution infrastructureable to cover large distances.

Semi-central facilities present interme-diate H2-production capacity. They pres-ent reduced distribution costs, since theyare sited closer to H

2 points of use.

Distributed facilities are small facili-ties located close to or at the pointof H2 use, reducing delivery costs. These facilities may present produc-tion capacities fitted to local demand. They require less investment than theother facilities, although unit produc-tion costs may be higher. n


Editor’s Note: The content for this column is supplied by In-tratec Solutions LLC (Houston; www.intratec.us) and edited byChemical Engineering. The analyses and models presented are

prepared on the basis of publicly available and non-confidentialinformation. The content represents the opinions of Intratec only.More information about the methodology for preparing analysiscan be found, along with terms of use, at www.intratec.us/che.

FIGURE 2. There are three types of hydrogen production facilities, and they differ in location and scale ofproduction

FIGURE 1. Steam methane reforming process for hydrogen production

1) Hydrotreater2) Desulfurizer3) Reformer reactor4) Water-gas shift

reactor5) Pressure-swing

adsorber system

ST SteamFU FuelBFW Boiler feed water

Naturalgas

H2 ST

ST

Air

FU BFW

H2product

1

2

4

5

3

80–480 km 40–160 km

Centralfacility

Distributedfacility

Semi-centralfacility

Hydrogenpoint of

use



Technology Profile

CHEMICAL ENGINEERING WWW.CHEMENGONLINE.COM JUNE 201534

Bio-based Methionine via Fermentation of Glucose


M

ethionine is an essentialamino acid that is not pro-duced by animals. Me-thionine must be obtained

from dietary sources, and its majoruse is as an animal nutritional-feedadditive. Methionine is mainly derivedfrom petrochemical sources. However,biochemical routes to obtain it havebeen researched. For instance, a jointventure between CJ Bio (cjbio.net) and Arkema (arkema.com) initiated methi-onine production using both renewableand petrochemical raw materials; andMetabolic Explorer (metabolic-explor-er.com), in partnership with Roquette

(roquette.com), developed a processfor 100% bio-based methionine.

The process

The production of L-methionine via amethod similar to the one proposed inpatents issued to Metabolic Explorerand Roquette is described below andpresented in Figure 1. In this process,L-methionine is produced through afed-batch aerobic fermentation, us-ing glucose as the carbon source and

ammonium thiosulfate as the nitrogenand sulfur source. Products obtainedfrom this process are L-methioninepowder (final product) and liquid L-methionine compound (byproduct). Media preparation. Culture mediaused in the fermentation are preparedby mixing water, corn syrup (70 wt.%glucose) and aqueous ammoniumthiosulfate solution in medium vessels. Fermentation. A recombinant mi-croorganism produces L-methioninethrough an aerobic fermentation pro-

cess, which is divided into two phas-es: batch and fed-batch. Fermentersare filled with the batch medium, andfermentation occurs until glucose ex-haustion, when the fed-batch phase isstarted by adding the fed-batch me-dium until the end of the fermentation.Clarification and demineralization. Culture broth is clarified by the re-moval of insoluble and soluble organicimpurities by centrifugation and ultra-filtration, respectively. Next, the brothis demineralized using ion-exchangeresins to remove cations and anions. Pre-concentration, crystallization

and filtration. L-methionine is con-

centrated in an evaporator prior tocrystallization under vacuum by waterevaporation. Then, precipitated methi-onine is sent to a vacuum filter, whereit is separated from the crystallizationmother liquor, which still contains dis-solved methionine. Products recovery. Filtered L-meth-ionine is directed to a dryer, where L-methionine powder with 99.5 wt.% ofdry matter (90 wt.% of methionine) isobtained. Evaporated water is sent to

a wash column before being used inmethionine washing during filtration. Also, the crystallization mother liquor

is acidified with HCl, then concentrat-ed, resulting in the liquid L-methioninebyproduct, which contains 60 wt.% ofdry matter (22 wt.% of methionine).

Economic performance An economic evaluation of the pro-cess was conducted based on datafrom Q1 2014. The following assump-tions were taken into consideration:

A facility erected on the U.S. Gulf•

Coast with a nominal capacity of87,600 ton/yr of L-methionine pow-

der and 86,700 ton/yr of liquidStorage autonomy of 15 d for feed-•

stock and 20 d for productsEstimated capital investment (total

fixed investment, working capital andinitial expenses) is about $430 million,and operating expenses are about$3,900/ton of L-methionine powder.

Global perspectiveHistorical methionine-consumption ispresented in Figure 2. Globally, West-

ern Europe, North America and North Asia are the major consumer regions.While most regions had similar growthin methionine consumption (averagegrowth rates around 5%/yr), North Asiastands out at 11%/yr. n

Edited by Scott Jenkins Editor’s Note: The content for this column is supplied byIntratec Solutions LLC (Houston; www.intratec.us) and editedby Chemical Engineering . The analyses and models presentedare prepared on the basis of publicly available and non-confidential information. The content represents the opinionsof Intratec only. More information about the methodology forpreparing analysis can be found, along with terms of use, atwww.intratec.us/che.

1,000

900

800700

600

500

400

300

200

100

02010 2011 2012 2013

n Africa n Central & South America n MiddleEast n North Asia n North America n South Asia &Oceania n West Europe

1) Batch medium vessels2) Fed-batch medium

vessels3) Fermenters4) Broth centrifuges5) Ultrafiltration unit6) Ion-exchange unit7) Evaporator8) Vacuum crystallizer9) Vacuum filter10) Vacuum dryer11) Vacuum wash column12) Liquid methionine

concentrator13) Cooling tower

14) Steam boilerCW Cooling waterST Steam

STCW

Corn syrup

1

2

5

H20

Ammoniumthiosulfate

solution

Corn syrup

H20

Ammoniumthiosulfate

solution

Air

Waterpurge

CW

CW

Insolubleimpurities

Soluble

impurities

Waterpurge

CW

ST

HCI

H20

L-methionineliquid compound

L-methioninepowder

To flare

CW

ST

3

4

6

7

8

9

10

12

11

13

14FIGURE 1. Bio-based methionine production via the direct fermentation of glucose

FIGURE 2. Methionine consumption by region



Technology Profile

CHEMICAL ENGINEERING WWW.CHEMENGONLINE.COM JULY 201533

Green FDCA Production


C

oncerns over limited sup-plies of petroleum-basedfeedstocks, coupled withthe global issue of climate

change, have increased demand fromconsumers and companies for moresustainable products. In this economicenvironment, chemical and polymerproduction from renewable resourcesis increasingly viewed as an attractivearea for investment.

Following this global trend, researchon the production of 2,5-furandicar-boxyllic acid (FDCA) has intensifiedover the last few years. FDCA hasthe potential to replace fossil-based

purified terephthalic acid (PTA) in theproduction of polyesters and otherpolymers containing an aromatic ring. The most promising application seg-ment for FDCA is for the productionof polyethylene furanoate (PEF). Thispolymer exhibits physical propertiesand applications similar to polyethyleneterephthalate (PET), and is producedby the use of FDCA, instead of PTA, inthe PET-production process.

Production processFigure 1 depicts an FDCA productionprocess from a glucose solution via thefuran pathway in a process similar tothat developed by Avantium Technolo-gies B.V. (Amsterdam, the Netherlands;www.avantium.com). This process canbe divided into three main areas: furansproduction, methoxy methyl furan (MMF)purification, and FDCA production.Furans production. The glucose con-tent of the feed is first enzymaticallyisomerized to fructose, and the effluent

is passed through ion-exchange resinsto separate unreacted glucose fromthe fructose. The former is recycled tothe isomerization reactor and the latteris sent to the furans conversion reac-tor. In the furans conversion reactor,crystallized fructose is solubilized witha 95 wt.% methanol aqueous solutionand sulfuric acid. The reaction is con-ducted at 50 bars and 200ºC. Fruc-tose is dehydrated to hydroxyl methylfuran (HMF) and, due to the excess ofmethanol, most HMF is converted toMMF, which is more stable than HMF.MMF purification. The reactor effluentis then sent to a distillation system to

recover the solvent and separate theMMF from other impurities. Methanoland water are recovered and returnedto the furans conversion area, whilemethyl levulinate is recovered as a by-product. Also, a heavies stream is re-covered to be used as boiler fuel, andfinally, purified MMF is recovered anddirected to the FDCA production area.FDCA production. Purified MMF ismixed with acetic acid solvent anda catalyst in the oxidation reactor to

be converted to FDCA. This stream isfiltered to recover crude FDCA. Theacetic-acid-rich liquid is purified us-ing two distillation columns. Catalystis also recovered, but it must be re-activated before being recycled to theprocess. Crude FDCA is mixed withwater and hydrogen in a hydrogena-tion reactor, where impurities are re-moved. The effluent from the reactoris crystallized, and purified FDCA is re-covered from the solid materials. Theliquid effluent is a waste stream.

Economic performance An economic evaluation of an FDCAplant was conducted, assuming afacility with a nominal capacity of300,000 ton/yr of purified FDCA con-structed on the U.S. Gulf Coast. In-cluded was a storage capacity equalto 20 days of operation for feedstocksand products.

Estimated capital expenses (totalfixed investment, working capital andinitial expenses) to construct the plantare about $600 million, while the oper-ating expenses are estimated at about$1,550/ton of purified FDCA.

Global perspective The main production cost, as in manycommodities-production processes,is the purchase of raw materials. Notonly is glucose syrup an expensivesource of glucose for this process,but also the conversion of fructose toMMF is relatively low, increasing theraw-material consumption. It wouldbe beneficial for the profitability of theprocess to integrate it with biomasstreatment, in order to obtain raw ma-

terials at lower costs. This processcould easily integrate with a biorefin-ery, where it is possible to achieve amore efficient valorization of the bio-mass used as feedstock. n


Editor’s Note: The content for this column is supplied byIntratec Solutions LLC (Houston; www.intratec.us) and editedby Chemical Engineering . The analyses and models presentedare prepared on the basis of publicly available and non-confidential information. The content represents the opinionsof Intratec only. More information about the methodology forpreparing analysis can be found, along with terms of use, atwww.intratec.us/che.

1) Isomerization reactor2) Evaporator3) Ion exchanger4) Crystallizers5) Furan conversion reactor6) MMF recovery system7) Oxidation reactor8) Rotary filter9) Hydrogenation reactor10) Crystallizers11) Centrifuges12) Rotating dryers13) Water removal column14) Acetic-acid-recovery column15) Cooling tower16) Steam boilerML Methyl levulinate byproductCW Cooling waterST Steam

Glucosesyrup

1

2

5

3

4

6

7

9

10

12

1314

CW

ST

Methanol sol.Sulfuric acid

Methanol

Waste water

Water forrecycle

Catalyst Acetic acid

Water

Hydrogen

Heavies to boiler

Waste water

Recoveredcatalyst

Waste water

Purified FDCA

CW CW CW CW

ST ST ST ST

ST

CW

ST

ST ST

8

15

16

11

ML

CW CW

FIGURE 1. FDCA production process via furans pathway similar to Avantium YXY process



Technology Profile

CHEMICAL ENGINEERING WWW.CHEMENGONLINE.COM AUGUST 201537

Lactic Acid Production via Glucose Fermentation


L

actic acid is a naturally occur-ring chiral acid most commonlyfound in the L(+) form. Drivenby a recent demand for more

sustainable products, the lactic acidmarket has expanded beyond thefood, pharmaceutical and cosmeticsindustries. Currently, a major growthapplication is for the production ofpolylactic acid (PLA) polymer. This bio-degradable polymer has the potentialto replace petrochemical-based plas-tics in applications such as packaging,paper coating, fibers, films and a widerange of molded plastic items.

The processFigure 1 depicts a process for lacticacid production via glucose fermen-tation similar to that developed byPurac (now Corbion; Amsterdam, theNetherlands; www.corbion.com). Thisprocess generates an 88 wt.% lac-tic acid solution in water that can beused as the raw material in polymer-ization processes. Rigorous control offermentation conditions and additionalpurification steps are necessary to en-

sure both optically and chemically purelactic acid. The quality of the PLA pro-duced from it depends on the purity.Fermentation. Glucose solution isfed to the fermenter. It undergoes acontinuous anaerobic bacterial fer-mentation, and is converted to lacticacid. The fermentation temperature iscontrolled by circulating cooling waterthrough internal coils, and the pH iscontrolled by the addition of CaCO3 solution, which reacts with lactic acidto form calcium lactate.

Recovery. The fermentation broth isdecanted and concentrated solids arefiltered, removing mostly biomass. Theaqueous solutions from the decanterand from the filter are acidified withsulfuric acid to convert calcium lactateinto soluble lactic acid and CaSO4,which precipitates. The slurry outletstream from the acidification vessel isfed to a belt filter for solids removal. The resulting aqueous filtrate is con-centrated in an evaporator, then fed toion-exchange columns for purification.Purification. Concentrated lactic acidis sent to a solvent-extraction system.First, it is contacted with an organic ex-

tractant solution, forming two phases:a lactic-acid-rich organic phase andan aqueous phase containing most ofthe contaminants. The organic phaseis then washed with water to removecontaminants. Next, it is contactedwith water at 80ºC to transfer the lacticacid into the aqueous phase.The re-sulting organic phase is washed withan NaOH solution and then with wa-ter to recover the extractant solution,while the lactic-acid-rich aqueous so-

lution is sent to an evaporator.The con-centrated lactic acid stream is fed to adistillation system where light ends areremoved. Heavy ends are hydrolyzedand recycled to the process. Lacticacid is recovered as a side stream.

Economic performance An economic evaluation for a lactic-acid production plant was conducted,assuming a facility with a nominal ca-pacity of 100,000 ton/yr of lactic acidsolution (88 wt.%) constructed on the

U.S. Gulf Coast, and including storage

capacity equal to 20 days of operationfor feedstock (no product storage). Es-timated capital expenses to constructthe plant are ~$200 million, while op-erating expenses are estimated at~$1,300/ton of lactic acid solution.

Global perspectiveMajor production-cost components forthis process are raw material acquisi-tion (mainly glucose and CaCO3 ) andCaSO4 byproduct disposal (Figure 2).

Recent improvement efforts focus onusing less costly cellulosic materialto obtain fermentable sugars, and ondeveloping low-pH-tolerant bacteriastrains, which would decrease bothCaCO3 use and CaSO4 generation. n


Editor’s Note: The content for this column is supplied byIntratec Solutions LLC (Houston; www.intratec.us) and editedby Chemical Engineering . The analyses and models presentedare prepared on the basis of publicly available and non-confidential information. The content represents the opinionsof Intratec only. More information about the methodology forpreparing analysis can be found, along with terms of use, atwww.intratec.us/che.

1. Fermenter2. Decanter3. Acidification vessel4. Rotary filter5. Belt filter6. Evaporator7. Ion exchange beds8. Solvent extraction system9. Evaporator10. Light end column11. Heavy end column12. Hydrolysis reactor13. Cooling tower14. Steam boiler


Glucose

CW

ST

Sodium hydroxide solutionSulfuric acid Process water

Extractantsolution

Lacticacid

Waste stream

CW

ST

ST

CW

STST

Calciumcarbonate

Biomass

Calciumsulfate

CW

Waste stream

Wastestream

12

5

3

4

67

9

10

12

13

14

8

11ST

ST

n Raw materialsn Calcium sulfate disposaln Utilitiesn Fixed costsn Depreciation

46%

11%

13%

13%

17%

FIGURE 2. Production costs breakdown

FIGURE 1. A lactic-acid production process from glucose, similar to the Purac process



Technology Profile

CHEMICAL ENGINEERING WWW.CHEMENGONLINE.COM SEPTEMBER 201544

Production of Polylactic Acid


P

olylactic acid (PLA) is the bio-degradable polymer of lacticacid, which is produced viafermentation processes. For

the last five decades, PLA has beenmainly used in biomedical applica-tions. The development of economi-cal production routes and a rising en-vironmental awareness by the generalpublic have pushed PLA uses to newmarkets, including consumer goodsand packaging applications.

High-molecular-weight PLA can beproduced through two main routes:an indirect route via a lactide interme-diate, and a direct polymerization by

polycondensation. The indirect routeis the most common in industry, andis employed by two of the major PLAproducers: NatureWorks LLC (Min-netonka, Minn.; www.natureworksllc.com) and Corbion-Purac (Amsterdam,the Netherlands; www.corbion.com).

In the indirect process, lactic acidis first oligomerized and depolymer-ized to produce lactide, a cyclic di-mer of lactic acid. Then, through aring-opening polymerization, lactide is

converted to polylactic acid.

PLA production processFigure 1 depicts a process for PLAproduction via a ring-opening polym-erization process similar to that devel-oped by NatureWorks. This processcan be divided into two main areas:oligomerization and lactide formation;and lactide polymerization. In this pro-cess, the optical properties of the lac-tic acid feedstock are preserved in thepolylactic acid.Oligomerization and lactide formation. Lactic acid is fed to the oli-

gomerization reactor, from whichwater vapor is removed, condensedand sent to a waste-treatment sys-tem. The resulting oligomer streamis mixed with catalyst and fed to thedepolymerization reactor, which is apacked column operating under vac-uum. Water, lactic acid and lactideare removed in the gaseous overheadstream. Most lactide and some im-purities are condensed and sent tothe drying column, where water is

removed from the crude lactide. Theremaining contaminants are removedin the purification column, leading to a99.9% pure lactide product.Lactide polymerization. The poly-mer-grade lactide undergoes a melt-phase polymerization. Molten lactide,a stabilizer and catalyst are fed to thepolymerization reactor. The polymermelt is then pumped to a devolatizeroperating under vacuum to removeunreacted lactide in the gaseous

stream, which is condensed and re-cycled to the purification column. Thedevolatilized melt polymer is extruded,pelletized and dried, resulting in PLAas a final product.

Economic performance An economic evaluation of a PLA plantwas conducted, assuming a facilitywith a nominal capacity of 100,000ton/yr of PLA constructed on the U.S.Gulf Coast. Storage capacity for thefinal product equal to 20 days of op-eration was assumed, but raw mate-rial storage was not included.

Estimated capital expenses (totalfixed investment, working capital andinitial expenses) to construct the plantare about $100 million, while the oper-ating expenses are estimated at about$1,900/ton of PLA.

Global perspective The recent decrease in oil prices hasreduced the production costs for mostfossil-fuel-based polymers, threaten-ing the competitiveness of PLA proj-

ects. Efforts to reduce PLA productioncosts must be focused on acquisitionof low-cost raw material. An alterna-tive to allow access to less costly rawmaterial would be the integration ofPLA production with lactic acid pro-duction from a renewable feedstock,such as biomass (Figure 2; Chem.Eng. August 2015, p. 37).

In an attempt to reach new mar-kets, some companies are research-ing methods for producing PLA with

enhanced physical properties. Thesemethods include the production of apolymer that consists of a mixture ofpoly(L-lactic acid) and poly(D-lacticacid). This polymer exhibits improvedphysical properties when comparedto conventional PLA, produced exclu-sively from L-lactic acid feedstock. n

Edited by Scott JenkinsEditor’s Note: The content for this column is supplied byIntratec Solutions LLC (Houston; www.intratec.us) and editedby Chemical Engineering. The analyses and models presentedare prepared on the basis of publicly available and non-confidential information. The content represents the opinions

of Intratec only. More information about the methodology forpreparing analysis can be found, along with terms of use, atwww.intratec.us/che.

1. Oligomerization reactor2. Depolymerization reactor3. Drying column4. Purification column5. Catalyst melting vessel6. Polymerization reactor7. Devolatizer8. Extruder & pelletizer9. Dryer10. Cooling tower11. Steam boiler


CW

ST

Air

CW

CW

ST

Lactic acid

Waste stream

1

2

5

3 4

67 9

10

8

11

Waste stream

Waste stream

Waste stream

Catalyst

Stabilizer

Catalyst

CWCW

CW CW

STST

Polylactic acid

FIGURE 1. The PLA production process occurs via a ring-opening polymerization process similar to that of NatureWorks

Polylac-tic acid

Solid waste Calcium sulfate

BiomassBiomass

treatment &hydrolysis

Glucosefermentationto lactic acid

Lactic acidconversionto lactide

Lactidering-opening

polymerization

FIGURE 2. PLA isproduced fromlactide rings



Technology Profile

CHEMICAL ENGINEERING WWW.CHEMENGONLINE.COM OCTOBER 201544

Ethylene Glycol Production


M

onoethylene glycol (MEG),also known as ethyleneglycol (EG) or simply gly-col, is a diol mostly used

for the production of polyester fibersand polyethylene terephthalate (PET)resins. It is also used in antifreezeapplications and in pharmaceuticalsand cosmetics. MEG is convention-ally produced through the hydrolysisof ethylene oxide (EO), which itself isobtained via ethylene oxidation.

The processFigure 1 depicts MEG productionfrom ethylene via a process similar to

the OMEGA catalytic process fromShell Global Solutions (The Hague,the Netherlands; www.shell.com). Inthe described process, MEG is pro-duced via EO, which is manufacturedin an integrated plant utilizing Shell EOtechnology. An important feature ofthe process is the negligible produc-tion of diethylene glycol (DEG) andtriethylene glycol (TEG), which occuras byproducts in other ethylene glycolproduction processes.

Ethylene oxide production. Ethyleneand oxygen are fed to a multi-tubularreactor, forming EO. This exothermicreaction, conducted in fixed beds in thereactor tubes, occurs in the gaseousphase with the use of a silver catalystsupported on alumina. Steam is gener-ated by the heat of reaction.Ethylene oxide recovery. The reac-tor product stream is fed to the EOabsorber for lights removal by waterquenching. Part of this gaseous over-head stream is recycled to the reactor,while the other part is sent to a carbon-

dioxide-removal unit composed of anabsorber and a stripper. In this unit,CO2 is separated to be used in ethyl-ene carbonate production.

A diluted EO stream removed fromthe absorber is fed to the EO stripper,where it is concentrated and recov-ered in the overheads. The crude EOstream is condensed. Residual light

gases are recovered from it and re-cycled to the reactor. The resulting EOstream is directed to the next section.Ethylene glycol production and purification. Ethylene oxide is reactedwith CO2, forming ethylene carbon-ate, which is then hydrolyzed to formMEG and CO2. Both reactions arecarried out in the liquid phase usinghomogeneous catalysts.

CO2 streams from the reaction stepsare recycled to the ethylene carbonate

reactor. MEG is purified in two distilla-tion columns where water is removed,leading to the final MEG product. Thecatalyst is separated and recycled tothe ethylene carbonate reactors.

Economic performance An economic evaluation of the pro-cess was conducted based on datafrom the first quarter of 2015, assum-ing a facility with a nominal capacityof 750,000 ton/yr of MEG constructedon the U.S. Gulf Coast.

Estimated capital expenses (total

fixed investment, working capital andinitial expenses) to construct the plantare about $630 million, while the oper-ating expenses are estimated at about$620/ton of MEG.

Global perspectiveShell OMEGA is the first process toenable ethylene glycol production via

a fully catalytic process. Accordingto the licenser, the process is able toachieve EO-to-EG converstion andselectivity near 100%, leading to pro-duction of MEG only.

However, although 40% of world-wide ethylene glycol production isderived from processes using Shelltechnologies, only three plants in theworld use the new Shell OMEGA tech-nology (Figure 2). As can be seen, twoof these plants are located in Asia,

which is the region of the world re-sponsible by the major share of MEGglobal consumption. China alone ac-counts for about 45% of global de-mand for MEG. n


Editor’s Note: The content for this column is supplied byIntratec Solutions LLC (Houston; www.intratec.us) and editedby Chemical Engineering. The analyses and models presentedare prepared on the basis of publicly available and non-confidential information. The content represents the opinionsof Intratec only. More information about the methodology for

preparing analysis can be found, along with terms of use, atwww.intratec.us/che.

FIGURE 2. PLANTS USING SHELL OMEGA TECHNOLOGY

Company Location Capacity (1,000 ton/yr) Start-up year

Lotte Chemical Daesan, South Korea 400 2008

Petro Rabigh Rabigh, Saudi Arabia 600 2009

Shell Jurong Island, Singapore 750 2009

1. Ethylene oxide reactor2. Ethylene oxide absorber3. Ethylene oxide stripper4. Lights removal unit5. CO2 absorber6. CO2 stripper7. Ethylene carbonate

reactors8. Hydrolysis reactors9. Water removal column10. MEG purification column11. Cooling tower12. Steam boiler

CW Cooling waterST SteamCW

ST

ST

1

7

9

10

12

8

11Waste stream

Steam

CW

ST

ST ST

Ethylene

O2

Ethylene oxide Catalyst recycle

Water

CO2 vent

CO2

MEG

CO2

2 3 4 5 6

FIGURE 1. Monoethylene glycol (MEG) production, according to a process similiar to the Shell OMEGA process



Technology ProfileEthylene Production via Cracking of Ethane-Propane


E

thylene is a critical buildingblock for the petrochemi-cal industry, and is amongthe most produced organic

compounds. It is usually produced insteam-cracking units from a range ofpetroleum-based feedstocks, such asnaphtha, and is used in the manufac-ture of several major derivatives.

The process The process shown in Figure 1 is asteam-cracking process for ethyleneproduction from an ethane-propanemixture. The process can be dividedinto three main parts: cracking and

quenching; compression and drying;and separation.Cracking and quenching. Initially, anethane-propane mixture is fed to fur-naces in which, under high-severityconditions, it is cracked, forming ethyl-ene, propylene and other byproducts. The furnace outlet stream is subse-quently fed to a water-based quench,to prevent further reactions and forma-tion of undesirable byproducts.

From a decanter downstream from

the quench tower, heavies, con-densed dilution steam, tar and cokeare removed. Cracked gas from thequench is then directed to compres-sion and separation.Compression and drying. The com-pression of the cracked gas is per-formed across five stages. After thethird stage of compression, carbon di-oxide and sulfur are removed from thecracked gas by caustic soda and wa-ter washes in a caustic scrubber. Thecompressed cracked gas is cooledand subsequently dried by molecularsieves that remove most of the water.

Separation. The dried cracked gasis fed to a cold box for the removalof hydrogen and light hydrocarbons,while minimizing ethylene losses.

At this point, condensates from thechilling train are fed to a series of sep-aration columns. In the first column(demethanizer), methane is obtainedfrom the top and further used in thecold box, while the bottom stream isfed to a second column (deethanizer).

The top of the deethanizer, com-posed primarily of ethylene and eth-ane, is fed to an acetylene converterand then fractionated in the C2-split-ter. In this column, lights are removed

from the overheads and recycled tothe compression system, while poly-mer-grade (PG) ethylene is drawn fromthe column as a side stream. Ethane,from C2-splitter bottoms, is recycledto the cracking furnaces.

The deethanizer bottom stream isfed to a depropanizer, which distillsC3 components in the overheads. This overhead stream is catalyticallyhydrotreated for methyl acetylene andpropadiene removal, and then fed to

the C3-splitter. In this column, lightsare removed from the overheadsand recycled to the compressors,while polymer-grade (PG) propyleneis drawn from the column as a sidestream. Propane from C3-splitter bot-toms is recycled to the cracking fur-naces. A C4+ stream is obtained fromthe depropanizer bottoms.

Economic performance An economic evaluation of the pro-cess was conducted based on datafrom the first quarter of 2015, consid-ering a facility with a nominal capacity

of 1,700,000 ton/yr of ethylene con-structed on the U.S. Gulf Coast.

Estimated capital expenses (totalfixed investment, working capital andinitial expenses) to construct the plantare about $2.37 billion, while the oper-ating expenses are estimated at about$360 per ton of ethylene produced.

Global perspectiveWith a global nominal capacity ofabout 155 million ton/yr, ethylene is

among the major petrochemicals pro-duced worldwide. The major part ofethylene production is consumed inthe manufacture of polyethylene, butethylene is also applied in the produc-tion of ethylene oxide, ethylene dichlo-ride and ethylbenzene (Figure 2). n


Editor’s Note: The content for this column is supplied byIntratec Solutions LLC (Houston; www.intratec.us) and editedby Chemical Engineering. The analyses and models presentedare prepared on the basis of publicly available and non-confidential information. The content represents the opinions

of Intratec only. More information about the methodology forpreparing analysis can be found, along with terms of use, atwww.intratec.us/che.

n Polyethylenen Ethylene oxiden Ethylene dichloriden Ethylbenzenen Others

59%

13%

13%

7%

8%

FIGURE 2. Ethylene is made into a host of products

1. Furnaces2. Quench column3. Compression system4. Caustic scrubber5. Drying unit6. Cold box7. Demethanizer column8. Deethanizer column9. C2-splitter10. Depropanizer column11. C3-splitter12. Cooling tower

Dilutionstream

CW2

3 3

4

5

Ethane/ propane

Tar andcoke todisposal

Water to wastetreatment

Lights tofuel

To wastetreatment

PGethyleneproduct

PGpropyleneb d t

1

6

7

8 9 10 11

Technology Profiles (Chemical Engineering)

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