Thermo Properties of Crude Oil? 1

[CoopEngineerVan]

Hi, I work in the pump industry and I'm trying to run some calculations for the design of a crude bottoms pump.

The problem is I'm having a really hard time finding thermodynamic information on crude bottoms.I have a bunch of hydrocarbon information books and sheets but all of them seem to be for the lighter fuels.

Specifically, I'll need the heat capacity, and saturation temperature of crude bottoms at 2.67bar and 2.42bar, but just the heat capacity or boiling point at 1atm would help a lot.

If anyone could point me in the direction of a book that would definitely have this information that I could order, or an online location with this kind of information, I would greatly appreciate it!

 

[mbeychok]

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Spikecomix:

First, let me say that I find it difficult to understand why you need the thermodynamic properties of "crude bottoms" to design a pump. I can understand why you might need viscosity and specific gravity information, but not why you need to know the "heat capacity" and the boiling points at 2.42 bar and 2.67 bar. By heat capacity, I assume that you mean the "specific heat capacity", Cp.

By "crude bottoms", do you mean the bottoms product from the atmospheric crude oil distillation tower in a petroleum refinery (i.e., what is usually called the "atmospheric residuum")? Or do you mean the bottoms from the subsequent vacuum distillation tower (i.e., what is usually called the "vacuum residuum")?

In either case, the boiling points of the residuums at 1 atmosphere absolute pressure will vary with (a) the design of the specific distilliation towers in any specific refinery and (b) with the geographic origin of the crude oil being processed in any specific refinery. Thus, you can expect different refineries to produce residuums with different atmospheric boiling points.

Having said that, if you are talking about an atmospheric residuum, the boiling point at 1 atmosphere absolute pressure might range from about 700 degrees Fahrenheit to 800 degrees Fahrenheit. Th specific heat capacity, Cp, of the atmospheric residuum might range from 0.7 to 0.8 Btu/lb/[°]F for a liquid temperature range of 700 to 800 degrees Fahrenheit.

If you are talking about a vacuum residuum, the atmospheric boiling point might be in excess of 900 degrees Fahrenheit. The specific heat capacity, Cp, of the vacuum residuum might range from 0.65 to 0.75 Btu/lb/[°]F for a liquid temperature of 800 to 1000 degrees F.

Milton Beychok
(Contact me at www.air-dispersion.com)
.

 

[CoopEngineerVan]

I'm trying to get the value of minimum flow based on the thermodynamic conditions of the crude bottoms it's pumping.

As far as I know, this pump will be pumping out the crude that cannot be distilled and that's the liquid medium I have to use for my calculations. I realize it won't be specific because of the differences in product-per-location, but right now I have zero information on the liquid and I need a way to make an estimate.

The problem is that if we don't have enough head, the temperature rise due to the mechanics of the pump could cause it to go above its (unknown) boiling point for that pressure, causing cavitation damage to the pump.

 

[mbeychok]

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Spikecomix:

this pump will be pumping out the crude that cannot be distilled

As pointed out in my previous response, the vacuum residuum will have a higher boiling point than the atmospheric residuum. So you must really ask your customer which residuum you are designing for. Simply saying "the crude that cannot be distilled" just doesn't get it done!!

Frankly, I may be wrong but off the top of my head, I very seriously doubt that your pump will cause either residuum to vaporize to any significant amount.



Milton Beychok
(Contact me at www.air-dispersion.com)
.

 

[TD2K]

Temperature rises across a pump, unless it's very high heads and/or very low efficiencies are a few degrees and usually much less. The highest I remember seeing was about 15F on a small butane pump (the pump motor was frankly doing a better job as a heater than providing head to the liquid but the pump efficiency in that case was really low).

The rise in pressure through the pump more than compensates for the increase in vapor pressure. Now, if you take a bubblepoint liquid, run it through a pump and then recycle part of that flow directly back to the pump suction you can have problems with some small amounts of vaporization. That's one reason it's preferred to recycle a minimum flow stream back to the suction vessel rather than just directly to pump suction.

 

[25362]

Distillation bottoms at equilibrium with vapors are provided with sufficient 'static head' for pump suctions, and/or quenched down to an acceptable temperature level to avoid thermal cracking, as in vacuum towers.

These procedures should help to overcome vaporization (NPSH) problems in the pump. Try to verify these "suction" conditions are in fact there.

 

[UmeshMathur]

This is several months after your original post, so I apologize.

For calculating the available pump NPSH, in this situation, you really don't need the vapor pressure at all, just the liquid head above pump centerline minus the suction line friction losses. This is because the liquid is at its bubble point.

Recall that all liquid draws from a distillation column are at their bubble point. Therefore, if you insist on calculating vapor pressure, an approximate value can be found from the sum of the following two terms:

(1) PHC = Ptop*(1 - yHC)
where PHC = partial pressure of the condensable portion of the overhead hydrocarbon vapor,
Ptop = top tower pressure, and
yHC = mole fraction in the tower overhead of the non-condensable gases, including steam.

yHC is generally available directly from a process simulation model of the tower but is not hard to estimate directly,

(2) Tower pressure drop, dP.

Thus VP = PHC + dP

Other properties, like fluid viscosity, can be estimated from the oil specific gravity and Watson characterization factor. Note, however, that viscosities of boiling hydrocarbons and petroleum fractions are generally in the range of 0.2 to 0.5 cS, low enough to ensure fully developed turbulent flow in piping. For this reason, in most situations, you can estimate friction factor independently of the Reynolds number.