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Single-cycle mixed-fluid LNG (PRICO) process

Part I: Optimal design

Sigurd Skogestad & Jørgen Bauck Jensen

Quatar, January 2009

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Single-cycle mixed fluid LNG (PRICO) process

Natural gas:• 45 kg/s (1.3 MTPA)• Feed at 40 bar and 30 °C

– 89.7 mol% C1, 5.5% C2, 1.8% C3, 0.1% C4, 2.8% N2

• Cooled to ~ -156 °C• Expansion to ~ 1 bar

– Flash gas may be used as fuel

• Liquefied natural gas (LNG) product at -162C

-162 °C

1 bar

45 kg/s

30 °C

40 bar

-156 °C

35 bar

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Single-cycle mixed fluid LNG (PRICO) process

Refrigerant:• Mixed fluid: ~ 33mol% C1, 35%

C2, 25% C4, 7% N2

• Partly condensed with sea water to ~ 30 °C

• Cooled to ~ -156 °C• Expansion to ~ 4 bar• Evaporates in NG HX• Super-heated ~ 10 °C• Compressed to ~ 22 bar

4 bar

Sup 10 °C

-156°C

19 bar

22 bar

45 kg/s

30 °C

40 bar

475 kg/s

30 °C

22 bar

-156 °C

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Compressor:• Max. pressure: 22 bar / 30 bar• Max. compressor suction

volume*: 317000 m3/h• Max. compressor head*: 263.6

kJ/kgOr: Max. compressor ratio* Pr, e.g. 5.5 (Price)

4. Max. compressor work: 77.5 MW / 120 MW

5. Minimum superheating: 10C

Design constraints

-162 °C

30 °C

40 bar

1 bar

3.33 kg/s * Design constraint only

30 °C

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Optimal design: TAC

min JTAC = Joperation + Jcapital

subject to c ≤ 0

• Joperation [$/year] is the annual operating cost– Joperation = Jutility + Jfeeds + Jproducts

• Jcapital [$/year] is the annualized cost of the equipment

• Total annualized cost (TAC) is minimized with respect to the design variables

– Flowsheet structure – Areas, sizes– Operating parameters (pressures etc.)

• Requires mixed integer non-linear programming• Our case Fixed structure Try a simpler approach

Maximize total profit = Minimize Total Annualized Cost (TAC):

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Idea: Specify ΔTmin to balance between

• operating costs (favoured by a low value)

• capital costs (favoured by a high value)

Simpler approach: Specify ΔTmin

-162 °C

30 °C

40 bar

1 bar

3.33 kg/s * Design constraint only

ΔTmin=2C*

30 °C

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Simple ΔTmin-method (Approach 1)

• ΔTmin (=2C) is added as an extra design constraint + minimize compressor work (Ws)

• BUT: The resulting design parameters (pressure etc.) are not optimal for the resulting process! – Reoptimizing reduces ΔTmin to about 1C and reduces work by about

5% (!)

– Cannot be fixed by iterating on ΔTmin

• Therefore: Approach 1 NOT USED

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Simplified TAC (sTAC)

Capital cost

Jcapital = Σi (Cfixed,i + Cvariable,i·Sini) / T

T – capital depriciation time, e.g. 10 years

1. Structure of plant given Cfixed,i = 0

2. Main equipment: Heat exchangers and compressor

3. Scaling exponent• n = 1 for compressor (Can then combine operation and capital cost!)• n = 0.65 for heat exchangers

4. Cvariable,i = C0 for all heat exchangers

Approach 2: Adjust C0 to get ΔTmin = 2C

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sTAC – Optimization problemMinimize cost

Case I: Feedrate (NG) givenCase II: Feedrate free

Here: Consider Case II.Minimize cost=

”Max. single-train LNG feed”

3.33 kg/s

1 bar

30 °C

40 bar30 °C

-162 °C

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Resulting “Max feed” sTAC:

• Minimization with respect to design parameters (AHOT

and ANG) and operating parameters (pressures etc.)– ANG: NG / cold refrigerant

– AHOT: hot refrigerant / cold refrigerant

• Here: Adjust C0 to obtain ΔTmin = 2C

• Other constraints c: depend on specific case

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Case 1 – Price and Mortko (1983)• Data

– LNG outlet temperature (before expansion) = -144 °C– 77.5 MW compressor power

– Maximum Ph = 22 bar

– Maximum Pr = Ph/Pl = 5.5

• Differences / uncertainties– Pure methane – Neglected removal of heavy components– Pressure losses (especially important at low pressure, e.g. compressor

suction)– Heating of fuel gas produces some LNG “for free”

• 3.7 % higher production compared with Price & Mortko– 44.6 kg/s LNG production– Gives large amount of fuel gas (7.7 kg/s, ~230 MW)

• Want to limit fuel to 3.33 kg/s, ~100 MW

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Case 2 – Limited fuel flow• Limitation on fuel flow instead of outlet temperature

– Maximum 3.33 kg/s of fuel (7.7. kg/s in Case 1)

– Outlet temperature down from -144 °C to -156 °C to get sufficient cooling with less flash gas (fuel)

– Production (with Ws=77.5 MW and Pr=5.5) reduced by 6 % compared with case 1

• From 44.6 kg/s to 41.7 kg/s

3.33 kg/s

77.5 MW

-162C41.7 kg/s

-156C

22 bar

4 bar

45 kg/s30C

475 kg/s30C

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Case 3,4 – Super-heating• Wish to find the optimal degree of super-heating

– 10.0 °C super-heating used for all cases except 3 and 4– Case 3; 11.6 °C super-heating increases production by 0.8 % compared with

case 2– Case 4; 25.7 °C super-heating decreases production by 1.3 % compared with

case 3

• Optimum is very flat in terms of super-heating• Some super-heating is necessary to protect the compressor• Some super-heating is optimal due to

– Internal heat exchange in the main heat exchanger

• However, the heat transfer coefficient in the super-heating region is lower than in the evaporating region

– This has not been considered here– Will tend to reduce the optimal amount of super-heating

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Case 5 – No pressure constraint• We have removed the following constraints

– Maximum Ph = 22 bar– Maximum Pr = Ph/Pl = 5.5

• Ph is increased to 50.4 bar and Pr is increased to 22• LNG production is increased by 11 % (from case 2)• The high pressure ratio is not possible with a single compressor

casing– The compressor head is too high– Two compressors in series will do the job

• Higher head [kJ/kg] gives lower refrigerant flow– Cooling duty per kg of refrigerant closely related to head– Less heat transfer area is needed since less warm refrigerant needs cooling

• The cost of an additional compressor casing is at least partly offset by the decreased heat transfer area and increased production

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Case 6,7 – Real GE Compressor• GE MCL1800 series compressor

– Centrifugal compressor with 1800 mm casing diameter

– Maximum suction volume is 380 000 m3/h active constraint

– Maximum discharge pressure Ph = 30 bar active constraint

• Case 6 – 77.5 MW; Same production as case 5 – Compressor head is 216 kJ/kg which is feasible with a single

compressor casing

• Case 7 – 120 MW; 71.1 kg/s of LNG product– Compressor head is 162 kJ/kg which is feasible with a single

compressor casing

– Corresponds to 2.0 million tons per annum (MTPA) with 330 operating days per year

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Case 8 – Liquid turbines• Expansion in liquid turbines

– Takes the pressure down to 2 bar above the saturation pressure

– Avoid vapour in the turbines

– Possible with two phase turbines?

• Production increased by 6.6 % compared with case 7– 75.8 kg/s ~ 2.2 MTPA per train

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Production vs. feed pressure

• Results for case 8• Achievable feed

pressure depends on – Location of heavy extraction

• Up-front or integrated• Recompression after heavy

extraction

– Feed compressor?

• Complicates the optimization problem– Very important for production

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Conclusion

• The constraints on the compressor performance is very important for the maximum production design case

– Maximum compressor head– Maximum compressor shaft work– Maximum compressor suction volume

• The feed pressure is very important for the achievable production

– We have assumed a fixed feed pressure of 40 bar

• A large PRICO train of 2.2 MTPA is feasible with a single compressor casing

– 2.0 MTPA without liquid turbines

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Conclusion II

• All the results presented here are with a minimum approach temperature ΔTmin = 2.0 °C– This is achieved by adjusting C0 in the optimization problem

• An alternative is to find a reasonable C0 and the use the same value for all cases– These results are presented in the paper

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Additional material

1. Table with results for all cases

2. Table with results for the alternative design method with constant C0

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Fixed C0 for all cases